Processes and systems for reforming of methane and light hydrocarbons to liquid hydrocarbon fuels

ABSTRACT

Processes for converting methane and/or other hydrocarbons to synthesis gas (i.e., a gaseous mixture comprising H 2  and CO) are disclosed, in which at least a portion of the hydrocarbon(s) is reacted with CO 2 . At least a second portion of the methane may be reacted with H 2 O (steam), thereby improving overall thermodynamics of the process, in terms of reducing endothermicity (ΔH) and the required energy input, compared to “pure” dry reforming in which no H 2 O is present. Such dry reforming (reaction with CO 2  only) or CO 2 -steam reforming (reaction with both CO 2  and steam) processes are advantageously integrated with Fischer-Tropsch synthesis to yield liquid hydrocarbon fuels. Further integration may involve the use of a downstream finishing stage involving hydroisomerization to remove FT wax. Yet other integration options involve the use of combined CO 2 -steam reforming and FT synthesis stages (optionally with finishing) for producing liquid fuels from gas streams generated in a number of possible processes, including the hydropyrolysis of biomass.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a continuation of U.S. application Ser. No.16/928,096, filed Jul. 14, 2020, now allowed, which is a divisional ofU.S. application Ser. No. 15/813,814, filed Nov. 15, 2017, now U.S. Pat.No. 10,738,247, which is hereby incorporated by reference in itsentirety.

FIELD OF THE INVENTION

Aspects of the invention relate to reforming catalysts and processes forthe reforming of methane and/or other hydrocarbons to produce asynthesis gas product comprising H₂ and CO, with further downstreamconversion to liquid hydrocarbons.

DESCRIPTION OF RELATED ART

The ongoing search for alternatives to crude oil, for the production ofhydrocarbon fuels is increasingly driven by a number of factors. Theseinclude diminishing petroleum reserves, higher anticipated energydemands, and heightened concerns over greenhouse gas (GHG) emissionsfrom sources of non-renewable carbon. In view of its abundance innatural gas reserves, as well as in gas streams obtained from biologicalsources (biogas), methane has become the focus of a number of possibleroutes for providing liquid hydrocarbons. A key commercial process forconverting methane into fuels involves a first conversion step toproduce synthesis gas (syngas), followed by a second, downstreamFischer-Tropsch (FT) conversion step. In this second step, the synthesisgas containing a mixture of hydrogen (H₂) and carbon monoxide (CO) issubjected to successive cleavage of C—O bonds and formation of C—C bondswith the incorporation of hydrogen. This mechanism provides for theformation of hydrocarbons, and particularly straight-chain alkanes, witha distribution of molecular weights that can be controlled to someextent by varying the FT reaction conditions and catalyst properties.Such properties include pore size and other characteristics of thesupport material. The choice of catalyst can impact FT product yields inother respects. For example, iron-based FT catalysts tend to producemore oxygenates, whereas ruthenium as the active metal tends to produceexclusively paraffins.

With respect to the first conversion step, upstream of FT, knownprocesses for the production of syngas from methane include partialoxidation reforming and autothermal reforming (ATR), based on theexothermic oxidation of methane with oxygen. Steam methane reforming(SMR), in contrast, uses steam as the oxidizing agent, such that thethermodynamics are significantly different, not only because theproduction of steam itself can require an energy investment, but alsobecause reactions involving methane and water are endothermic. Morerecently, it has also been proposed to use carbon dioxide (CO₂) as theoxidizing agent for methane, such that the desired syngas is formed bythe reaction of carbon in its most oxidized form with carbon in its mostreduced form, according to:

CH₄+CO₂→2CO+2H₂.

This reaction has been termed the “dry reforming” of methane, andbecause it is highly endothermic, thermodynamics for the dry reformingof methane are less favorable compared to ATR or even SMR. However, thestoichiometric consumption of one mole of carbon dioxide per mole ofmethane has the potential to reduce the overall carbon footprint ofliquid fuel production, providing a “greener” consumption of methane.This CO₂ consumption rate per mole of feed increases in the case ofreforming higher hydrocarbons (e.g., C₂-C₆ paraffins), which may bedesired, for example, if hydrogen production (e.g., for refineryprocesses) is the objective. In any event, the thermodynamic barriernonetheless remains a major challenge and relates to the fact that CO₂is completely oxidized and very stable, such that significant energy isneeded for its activation as an oxidant. In view of this, a number ofcatalyst systems have been investigated for overcoming activation energybarrier for the dry reforming of methane, and these are summarized, forexample, in a review by Lavoie (FRONTIERS IN CHEMISTRY (November 2014),Vol. 2 (81): 1-17), identifying heterogeneous catalyst systems as beingthe most popular in terms of catalytic approaches for carrying out thisreaction.

Whereas nickel-based catalysts have shown effectiveness in terms oflowering the activation energy for the above dry reforming reaction, ahigh rate of carbon deposition (coking) of these catalysts has also beenreported in Lavoie. The undesired conversion of methane to elementalcarbon can proceed through methane cracking (CH₄→C+2H₂) or the Boudouardreaction (2CO→C+CO₂) at the reaction temperatures typically required forthe dry reforming of methane. Therefore, although this reaction has beeninvestigated as a promising route for syngas production, thecommercialization of this technology, unlike other reformingtechnologies such as ATR and SMR, remains unrealized. This is due inlarge part to high rates of carbon formation and the accompanyingdeactivation of catalysts through coking, as encountered in the use ofdry reforming catalyst systems that operate under conditions proposed todate. Finally, whereas other conventional reforming technologies haveproven to be economically viable, these processes, and particularly SMR,are known to require significant upstream capital and operating expensesfor the removal of sulfur and other poisons of the catalysts used.Otherwise, commercially acceptable periods of operation from a givencatalyst loading cannot be achieved. Satisfactory solutions to these andother problems relating to the conventional reforming of hydrocarbonsfor the production of syngas and/or hydrogen have been sought but notachieved.

SUMMARY OF THE INVENTION

Aspects of the invention are associated with the discovery of reformingcatalysts and processes for converting methane and/or other hydrocarbonsto synthesis gas (i.e., a gaseous mixture comprising H₂ and CO) byreacting at least a portion of such hydrocarbon(s) with CO₂. Preferably,according to a CO₂-steam reforming reaction, at least a second portionof the hydrocarbon(s) (e.g., comprising the same hydrocarbon(s) as inthe first portion) is reacted with H₂O (steam), thereby improvingoverall thermodynamics of the process, in terms of reducingendothermicity (ΔH) and the required energy input, compared to “pure”dry reforming in which no H₂O is present. Representative reformingcatalysts advantageously possess high activity and thereby can achievesignificant levels of hydrocarbon (e.g., methane) conversion attemperatures below those used conventionally for dry reforming. Thesehigh activity levels, optionally in conjunction with using H₂O toprovide at least a portion of the oxidant, contribute to an overalloperating environment whereby coke formation is reduced and usefulreforming catalyst life may be significantly extended.

Yet further important advantages reside in the sulfur tolerance ofreforming catalysts described herein, whereby a pretreatment of amethane-containing feedstock (e.g., natural gas), or otherhydrocarbon-containing feedstock, to reduce the concentration of H₂S andother sulfur-bearing contaminants is not required according to preferredembodiments, or is at least not as rigorous as in conventional reformingtechnologies. Also, to the extent that downstream sulfur removal may bedesirable, such as prior to an FT synthesis step, this may be greatlysimplified, considering that all or at least a substantial portion ofsulfur-bearing contaminants other than 1-H₂S, such as mercaptans, can beoxidized in a dry reforming or CO₂-steam reforming reaction as describedherein to SO₂, thereby rendering standard acid gas treatment (e.g.,scrubbing) as a suitable and relatively simple option for suchdownstream sulfur removal.

Overall, improvements associated with the processes and reformingcatalysts described herein are of commercial significance in terms ofrendering dry reforming processes, or otherwise CO₂ and steam reforming(i.e., “CO₂-steam reforming”) processes, as an economically viablealternative to conventional technologies such as autothernal reforming(ATR) and steam methane reforming (SMR). Moreover, the synthesis gasaccording to these processes may be produced with a favorable molarH₂:CO ratio (e.g., about 2:1) for downstream processing via theFischer-Tropsch (FT) reaction, or at least with a molar ratio that maybe readily adjusted to achieve such favorable values.

The demonstrated ability of CO₂-steam reforming processes describedherein to produce synthesis gas products with favorable molar H₂:COratios, in a stable manner and with tolerance to sulfur-bearingcontaminants that are often present in sources of methane (e.g., naturalgas) and other light hydrocarbons, provides advantages in the use ofthese processes with additional steps for producing liquid hydrocarbons,for example gasoline- and diesel boiling-range hydrocarbon fractions.These advantages include greater simplicity of overall liquidhydrocarbon production processes, which may, for example, require feweraddition, separation, and/or recycle steps compared to conventionalprocesses. This results not only in cost savings, but also in thepossibility of providing such overall processes in an easilytransportable (e.g., skid mounted) configuration, which may be broughtto sources of natural gas, or other sources of components of gaseousmixtures as described herein, from which sources the transport of suchcomponents to conventional brick and mortar production facilities wouldotherwise be problematic. Advantages also include increased flexibilityin terms of opportunities for integration with a wide variety ofprocesses that generate CO₂— and/or light hydrocarbon-containing gasstreams, including biomass conversion processes, fermentation processes,and industrial processes that generate CO₂-containing waste gases.

These and other embodiments, aspects, and advantages relating to thepresent invention are apparent from the following Detailed Description.

BRIEF DESCRIPTION OF THE DRAWINGS

A more complete understanding of the exemplary embodiments of thepresent invention and the advantages thereof may be acquired byreferring to the following description in consideration of theaccompanying figures, in which the same reference numbers are used toidentify the same or similar features.

FIGS. 1A and 1B depict flowschemes that illustrate representative dryreforming and CO₂-steam reforming processes as described herein.

FIG. 2 illustrates the relationship between pressure in aFischer-Tropsch (FT) reactor and the level of CO conversion obtained,with other operating conditions remaining constant.

FIG. 3 depicts a flowscheme in which a dry reforming or CO₂-steamreforming process, such as depicted in FIG. 1A or 1 i, is integratedwith downstream processing steps for producing liquid hydrocarbons.

FIG. 4 depicts a flowscheme in which a dry reforming or CO₂-steamreforming process, such as depicted in FIG. 1A or 1B, is used with aprocess for producing a renewable hydrocarbon fuel from thehydropyrolysis of biomass.

FIG. 5 depicts a flowscheme in which dry reforming or CO₂-steamreforming is integrated in an overall liquid hydrocarbon productionprocess, such as depicted in FIG. 3 , which is used with a process forproducing a renewable hydrocarbon fuel from the hydropyrolysis ofbiomass.

FIG. 6 depicts a flowscheme of a process for producing a renewablehydrocarbon fuel from the hydropyrolysis of biomass, such as a processwith which a dry reforming or CO₂-steam reforming process may be used,as depicted in FIG. 4 , or with which an overall liquid hydrocarbonproduction process may be integrated, as depicted in FIG. 5 .

FIG. 7 depicts a flowscheme in which dry reforming or CO₂-steamreforming is integrated in an overall liquid hydrocarbon productionprocess, such as illustrated in FIG. 3 , which is used in a hydrogenproduction process.

FIG. 8 illustrates the high activity, in terms of methane conversion, ofreforming catalysts as described herein.

FIG. 9 illustrates the relationship between the molar H₂:CO ratio of thesynthesis gas product and the molar H₂O/CO₂ ratio of the gaseous mixturein a CO₂-steam reforming reactor (as a combined feed) at differentreaction temperatures, in the case of representative CO₂-steam reformingprocesses.

FIGS. 10 and 11 illustrate the long term operational stability ofreforming catalysts as described herein, in a CO₂-steam reformingprocess over an extended operating period.

The figures should be understood to present illustrations of processesand certain associated results and parameters and/or principlesinvolved. In order to facilitate explanation and understanding, FIGS.1A, 1 i, 3-7, 10, and 11 provide a simplified overview, with theunderstanding that these figures and elements shown are not necessarilydrawn to scale. Valves, instrumentation, and other equipment and systemsnot essential to the understanding of the various aspects of theinvention are not shown. As is readily apparent to one of skill in theart having knowledge of the present disclosure, processes for convertinghydrocarbons such as methane, by dry reforming or CO₂-steam reforming,will have configurations and elements determined, in part, by theirspecific use.

DETAILED DESCRIPTION

The expressions “wt-%” and “mol-%,” are used herein to designate weightpercentages and molar percentages, respectively. The expressions“wt-ppm” and “mol-ppm” designate weight and molar parts per million,respectively. For ideal gases, “mol-%” and “mol-ppm” are equal topercentages by volume and parts per million by volume, respectively.

As used herein, terms such as “C₄ ⁺ hydrocarbons,” “C₂₀ ⁺ hydrocarbons,”“C₄-C₁₉ hydrocarbons,” etc. refer to hydrocarbons having greater than 4carbon atoms, hydrocarbons having greater than 20 carbon atoms,hydrocarbons having from 4 to 19 carbon atoms, etc., respectively.Unless otherwise stated, these terms do not imply that hydrocarbonshaving all carbon numbers according to the specified ranges mustnecessarily be present. Unless otherwise stated, e.g., by thedesignation “normal C₂₀ ⁺ hydrocarbons,” hydrocarbons of all types areincluded in such terms (e.g., normal, branched, aromatic, naphthenic,olefinic, etc.).

The term “gaseous mixture” refers to the mixture comprising at least ahydrocarbon such as methane and also comprising CO₂ as an oxidant, whichis subjected to dry reforming or CO₂-steam reforming (if water is alsopresent in the gaseous mixture) by contact with a reforming catalyst asdescribed herein. The term “gaseous mixture” refers generally to thismixture being completely or at least predominantly in the gas phaseunder conditions used for dry reforming or CO₂-steam reforming(“reforming conditions”), including the temperatures and pressuresdescribed herein as being suitable for these reactions. The term“gaseous mixture” does not preclude the presence of compounds in thismixture that, like water, are liquid under conditions of ambienttemperature and pressure. Such compounds can include hydrocarbons foundin liquid fuels including naphtha and jet fuels, for example C₆-C₁₆hydrocarbons.

The terms “naphtha boiling-range hydrocarbons” and “gasolineboiling-range hydrocarbons” refer to a hydrocarbon fraction comprisinghydrocarbons having boiling points within an initial (“front-end”)distillation temperature of 35° C. (95° F.), characteristic of C₅hydrocarbons, and an end point distillation temperature of 204° C. (399°F.). The term “jet fuel boiling-range hydrocarbons” refers to ahydrocarbon fraction comprising hydrocarbons having boiling pointswithin a front-end distillation temperature of 204° C. (399° F.) and anend point distillation temperature of 271° C. (520° F.). The term“diesel boiling-range hydrocarbons” refers to a hydrocarbon fractioncomprising hydrocarbons having boiling points within a front-enddistillation temperature of 204° C. (399° F.) and an end pointdistillation temperature of 344° C. (651° F.). Accordingly, “dieselboiling-range hydrocarbons” encompass “jet fuel boiling-rangehydrocarbons,” but also include “heavy diesel boiling-rangehydrocarbons” having boiling points within a front-end distillationtemperature of 271° C. (520° F.) and an end point distillationtemperature of 344° C. (651° F.). The term “VGO boiling-rangehydrocarbons” refers to a hydrocarbon fraction comprising hydrocarbonshaving boiling points within a front-end distillation temperature of344° C. (651° F.) and an end point distillation temperature of 538° C.(1000° F.). These front end and end point distillation temperatures ofhydrocarbon fractions, such as naphtha boiling-range hydrocarbons,gasoline boiling-range hydrocarbons, jet fuel boiling-rangehydrocarbons, and diesel boiling-range hydrocarbons, which are alsocharacteristic of respective petroleum derived naphtha, gasoline, jetfuel, and diesel boiling-range fractions, are determined according toASTM D86, with the end point being the 95% recovery value.

The term “substantially,” as used in the phrase “substantially same” or“substantially the same,” in reference to a given parameter, is meant toencompass values that deviate by less than 5% with respect to thatparameter when measured in absolute terms (e.g., absolute temperature orabsolute pressure). The term “substantially all” or “substantially allof” means “at least 95% of.” The term “substantially complete” means “atleast 95% complete.”

Embodiments of the invention are directed to a process for producing asynthesis gas product (syngas), the process comprising contacting agaseous mixture comprising (i) methane and/or other hydrocarbon(s)(e.g., any of CH₄, C₂H₆, C₂H₄, C₃H₈, C₃H₆, C₄H₁₀, C₄H₈, C₅H₁₂, C₅H₁₀,higher molecular weight hydrocarbons, and mixtures thereof) and (ii)CO₂, with a reforming catalyst comprising at least one (e.g., two, ormore than two) noble metals on a solid support comprising cerium oxide.It is possible that CO₂ alone can serve as the oxidant for the methaneand/or other hydrocarbon(s) to CO and H₂ according to the dry reformingof such hydrocarbons, which in the case of alkanes, for example, can begeneralized as:

C_(n)H_(2n+2) +nCO₂→2nCO+(n+1)H₂.

In preferred embodiments a combination of CO₂ and 1-120 can serve as theoxidant, that is, in embodiments in which the gaseous mixture furthercomprises H₂O. The reaction in this case is a “CO₂-steam reforming”reaction, which also includes steam reforming as a route for producingsyngas from methane and/or other hydrocarbons, which in the case ofalkanes, for example, can be generalized as:

C_(n)H_(2n+2) +nH₂O→nCO+(2n+1)H₂.

Whereas the theoretical molar H₂:CO ratio of a synthesis gas productformed from the dry reforming of methane is 1, the addition of steamreforming, in the CO₂-steam reforming of methane, advantageouslyprovides the potential to increase this molar ratio to values morefavorable for downstream Fischer-Tropsch synthesis to produce liquidhydrocarbons, according to the reaction:

(2n+1)H₂ +nCO→C_(n)H_(2n+2) +nH₂O.

From this, it can be observed that C₄ ⁺ hydrocarbons, such as C₄-C₁₂hydrocarbons, which are desirable as fuels or components of fuels, areformed ideally at molar H₂:CO ratios approaching 2. Importantly, the useof steam (H₂O) as an oxidant in combination with CO₂ provides anadvantageous “handle” or control parameter for adjusting the molar H₂:COratio of the synthesis gas product over a wide range of CO₂-steamreforming conditions. In fact, for any given set of such conditions(e.g., conditions within the CO₂-steam reforming reactor such astemperature, pressure, weight hourly space velocity, and reformingcatalyst formulation) under which the combined CO₂ and steam reformingreactions are carried out, a relationship can be established between themolar H₂O:CO₂ ratio of the gaseous mixture (e.g., combined CO₂-steamreforming reactor feed) and the molar H₂:CO ratio of the synthesis gasproduct (e.g., CO₂-steam reforming reactor effluent). Whereas the dryreforming and steam reforming of hydrocarbons other than methane produceH₂ and CO at other molar ratios, directionally the same shifts oradjustments in product yields may be achieved by varying the relativeamounts of the oxidants H₂O and CO₂ in the gaseous mixture that issubjected to CO₂-steam reforming. Accordingly, embodiments of theinvention are directed to a CO₂-steam reforming process comprisingdetermining a molar H₂:CO ratio of the synthesis gas product and, basedon the molar H₂:CO ratio, adjusting a molar H₂O:CO₂ ratio of the gaseousmixture toward a target molar H₂:CO ratio of the synthesis gas product,for example a target molar H₂:CO ratio of 2:1, or otherwise a targetmolar H₂:CO ratio range generally from about 1.5:1 to about 2.5:1,typically from about 1.5:1 to about 2.3:1, and often from about 1.8:1 toabout 2.2:1.

More specifically, the molar H₂O:CO₂ ratio of the gaseous mixture may beincreased to increase, toward the target molar H₂:CO ratio, an observedmolar H₂:CO ratio of the synthesis gas product that is below the target.Conversely, the molar H₂O:CO₂ ratio of the gaseous mixture may bedecreased to decrease, toward the target molar H₂:CO ratio, an observedmolar H₂:CO ratio of the synthesis gas product that is above the target.Any such adjustments to the molar H₂O:CO₂ ratio of the gaseous mixturemay be performed, for example, by adjusting the flow rate(s) of one ormore components of the gaseous mixture (e.g., combined feed), such asone or more of a methane-containing feedstock (or hydrocarbon-containingfeedstock generally), a CO₂-containing oxidant, and an H₂O-containingoxidant, relative to the flow rate(s) of one or more other of suchcomponents. According to a specific example, the molar H₂O:CO₂ ratio ofthe combined feed to the CO₂-steam reforming reactor may be increased ordecreased, by increasing or decreasing, respectively, the flow rate ofsteam (as the H₂O-containing oxidant), thereby resulting in a respectiveincrease or decrease in the molar H₂O:CO₂ ratio of the gaseous mixture.

In addition to providing the ability to control the molar H₂:CO ratio ofthe synthesis gas product over a favorable range of values, the use ofsteam (H₂O) as an oxidant in combination with CO₂ furthermoresurprisingly reduces the rate of carbon (coke) formation compared topure dry reforming, thereby extending the life of catalysts as describedherein. Accordingly, further embodiments of the invention are directedto a CO₂-steam reforming process in which the rate of carbon formation(e.g., using suitable ratios or concentrations/partial pressures of CO₂and H₂O oxidants, in combination with a reforming catalyst as describedherein) is less than the rate of carbon formation of a baseline process(i.e., baseline dry reforming process), in which all parameters aremaintained the same, except for the replacement of H₂O in the gaseousmixture (e.g., combined CO₂-steam reforming reactor feed) with anequimolar amount of oxygen as CO₂ (i.e., replacement of the moles of H₂Owith ½ the moles of CO₂). Coupled with this comparatively lower carbonformation relative to the baseline process, the synthesis gas productmay have a molar H₂/CO ratio as described herein (e.g., from about 1.5:1to about 2.3:1).

CO₂-steam reforming, as described herein, can be performed to produce asynthesis gas product having a favorable molar H₂:CO ratio in the rangesdescribed above, such as from about 1.5:1 to about 2.5:1, from about1.5:1 to about 2.3:1, and from about 1.8:1 to about 2.2:1. Such ranges,encompassing 2:1, are particularly advantageous in the case ofdownstream processing of the synthesis gas product in an FT synthesisstage, as described herein, to produce liquid hydrocarbons. Inparticular, a step of converting H₂ and CO in the synthesis gas productto hydrocarbons, including C₄ ⁺ hydrocarbons (including hydrocarbonsthat are liquid at ambient temperature and pressure) that are providedin an FT product, may be carried out with an FT feed having asubstantially same H₂:CO molar ratio as in the synthesis gas product,produced by the upstream CO₂-steam reforming. That is, the FT feed maybe obtained preferably without adjustment of the H₂:CO molar ratio ofthe synthesis gas product, such as by adding or removing H₂ and/or CO orotherwise converting or producing these components (e.g., without addingH₂ to increase this molar ratio and/or without the use of a separatewater-gas shift reaction or reverse water-gas shift reaction). Accordingto some embodiments, the FT feed may be obtained at substantially thesame H₂:CO molar ratio as in the synthesis gas product, by condensingwater from this product, prior to converting H₂ and CO to hydrocarbonsin the FT synthesis stage.

According to some embodiments, the FT feed may be obtained without anychange in composition of the synthesis gas product. For example, some orall of the synthesis gas product may be used directly in the FTsynthesis stage without any intervening operation that would impact itscomposition (e.g., by the addition, removal, or conversion of componentsthat would alter this composition).

The above ranges of molar H₂:CO ratios of the synthesis gas product,encompassing 2:1, are likewise advantageous in the case of downstreamprocessing of the synthesis gas product in a methanol production stageto produce methanol according to the reaction 2H₂+CO→CH₃OH. Inparticular, a step of converting H₂ and CO in the synthesis gas productto methanol that is provided in a methanol product, may be carried outwith a methanol synthesis feed having a substantially same H₂:CO molarratio as in the synthesis gas product, produced by the upstreamCO₂-steam reforming. That is, the methanol synthesis feed may beobtained preferably without adjustment of the H₂:CO molar ratio of thesynthesis gas product, such as by adding or removing H₂ and/or CO orotherwise converting or producing these components (e.g., without addingH₂ to increase this molar ratio and/or without the use of a separatewater-gas shift reaction or reverse water-gas shift reaction). Accordingto some embodiments, the methanol synthesis feed may be obtained atsubstantially the same H₂:CO molar ratio as in the synthesis gasproduct, by condensing water from this product. According to someembodiments, the methanol synthesis feed may be obtained without anychange in composition of the synthesis gas product. For example, some orall of the synthesis gas product may be used directly in the methanolproduction stage without any intervening operation that would impact itscomposition (e.g., by the addition, removal, or conversion of componentsthat would alter this composition). Methanol production from thesynthesis gas product may be carried out at a temperature from about204° C. (400° F.) to about 316° C. (600° F.) and a pressure from about4.5 MPa (650 psig) to about 11.7 MPa (1700 psig). Methanol synthesiscatalysts typically comprise Cu and ZnO, supported on a metal oxide suchas alumina (Al₂O₃).

In the case of production of methanol from the synthesis gas product,this methanol may be further reacted in a dehydration stage to producedimethyl ether (DME) according to the reaction 2CH₃OH→CH₃OCH₃+H₂O.Catalysts and conditions for conducting this reaction stage aredescribed, for example, in U.S. Pat. No. 5,037,511; US 2004/0034255; andU.S. Pat. No. 8,451,630. Alternatively, DME may be produced directlyfrom the synthesis gas product in a direct DME production stage, withoutan intervening methanol production stage. In this regard, dry reforming,as described herein, can be performed to produce a synthesis gas producthaving a favorable molar H₂:CO ratio in ranges encompassing 1:1 that aresuitable for carrying out the reaction 3H₂+3CO→CH₃O CH₃+CO₂, asdescribed, for example, in Takeishi et al. (Recent Advances in Energy &Environment). Suitable molar H₂:CO ratios are from about 0.5:1 to about1.5:1, from about 0.5:1 to about 1.3:1, or from about 0.8:1 to about1.2:1. In particular, a step of converting H₂ and CO in the synthesisgas product to DME that is provided in a DME product, may be carried outwith a DME synthesis feed having a substantially same H₂:CO molar ratioas in the synthesis gas product, produced by the upstream dry reforming.That is, the DME synthesis feed may be obtained preferably withoutadjustment of the molar H₂:CO ratio of the synthesis gas product, suchas by adding or removing H₂ and/or CO or otherwise converting orproducing these components (e.g., without adding H₂ to increase thismolar ratio and/or without the use of a separate water-gas shiftreaction or reverse water-gas shift reaction). According to someembodiments, the DME synthesis feed may be obtained at substantially thesame molar H₂:CO ratio as in the synthesis gas product, by condensingwater from this product. According to some embodiments, the DMEsynthesis feed may be obtained without any change in composition of thesynthesis gas product. For example, some or all of the synthesis gasproduct may be used directly in the direct DME production stage, withoutany intervening operation that would impact its composition (e.g., bythe addition, removal, or conversion of components that would alter thiscomposition).

In addition to producing a synthesis gas product having a desirablemolar H₂:CO ratio that can be tailored to particular, downstreamreaction steps as described above, reforming catalysts as describedherein furthermore exhibit a surprising degree of sulfur tolerance,which is particularly advantageous, for example, in the case ofmethane-containing feedstocks comprising or derived from natural gasthat, depending on its source, may contain a significant concentration(e.g., several weight percent by volume or more) of H₂S. In this regard,conventional steam methane reforming (SMR) processes requirepretreatment to reduce the feed total sulfur content to typically lessthan 1 mol-ppm to protect the reforming catalyst from sulfur poisoning.In contrast, according to representative embodiments of the presentinvention, the gaseous mixture or any of its components, particularlythe hydrocarbon-containing feedstock, is not subjected to, or otherwisehas not undergone, a sulfur removal pretreatment step. Such embodimentsprovide substantial economic benefits over known processes withstringent desulfurization requirements and associated expenses, asnecessary to achieve favorable reforming catalyst life. In contrast tosuch known processes, a gaseous mixture in a dry reforming or CO₂-steamreforming process as described herein may comprise sulfur generally atany concentration representative of the source of the hydrocarbonfeedstock, such as natural gas, not having undergone pretreatment forsulfur removal, but also accounting for the potential dilution of thesulfur when combined with other components of the gaseous mixture (e.g.,CO₂) having a lower sulfur concentration. For example, the gaseousmixture may comprise generally at least about 1 mole-ppm (e.g., fromabout 1 mol-ppm to about 10 mol-%) total sulfur (e.g., as H₂S and/orother sulfur-bearing contaminants). The gaseous mixture may comprisetypically at least about 10 mol-ppm (e.g., from about 10 mol-ppm toabout 1 mol-%) and often at least about 100 mol-ppm (e.g., from about100 mol-ppm to about 1000 mol-ppm) of total sulfur. For example, a rangefrom about 500 mol-ppm to about 1000 mol-ppm of total sulfur, accordingto particular embodiments, generally poses no, or at least a negligible,adverse effect on the stability of reforming catalysts as describedherein.

With respect to sulfur tolerance of reforming catalysts describedherein, further aspects of the invention are associated with thediscovery that higher levels (concentrations) of sulfur in the gaseousmixture may be compensated for by increasing the reaction temperature,i.e., temperature of the bed of reforming catalyst as described herein,contained in a reforming reactor (which may be either a dry reformingreactor or a CO₂-steam reforming reactor, with the latter term beingapplicable to the gaseous mixture within the reactor comprising both CO₂and H₂O). That is, increased sulfur concentrations have been found toimpact reforming catalyst activity, as measured by decreased conversionof methane and/or or other hydrocarbon(s) in the gaseous mixture, if allother operating parameters remain unchanged. However, the desiredconversion level can be restored by increasing the reaction temperature.For example, under certain operating conditions, a 28° C. (50° F.)increase can be sufficient to restore a loss in reforming catalystactivity that accompanies a concentration of 800 mol-ppm H₂S in thegaseous mixture, relative to the activity without any sulfur in thegaseous mixture. Accordingly, embodiments of the invention are directedto a dry reforming process or a CO₂-steam reforming process as describedherein comprising determining a conversion of methane and/or otherhydrocarbon(s) (e.g., a conversion of combined C₁-C₄ hydrocarbons orcombined C₁-C₃ hydrocarbons), or otherwise determining a sulfur level(such as an H₂S level) in the gaseous mixture or synthesis gas productand, based on the conversion or sulfur level, adjusting the reactiontemperature toward a target conversion of methane and/or otherhydrocarbon(s), for example a target conversion of at least about 75%(e.g., any specific conversion value in the range from about 75% toabout 100%), such as a target conversion of at least about 85% (e.g.,any specific conversion value in the range from about 85% to about 99%).

Importantly, however, such decreases in the activity of reformingcatalysts described herein, accompanying increases in the concentrationof sulfur in the gaseous mixture, are not further accompanied by anyappreciable loss in reforming catalyst stability. That is, thecompensating reforming reactor temperature increases, as describedherein to offset higher sulfur levels, do not significantly impact theability of the reforming catalyst to achieve stable operatingperformance with respect to dry reforming or CO₂-steam reforming over anextended period. This finding is contrary to expectations based onconventional reforming technologies, in which the presence of even smallquantities (e.g., mol-ppm levels) of sulfur in feeds must be preventedto avoid deactivation and costly premature replacement of the catalyst.A characteristic sulfur tolerance, or activity stability in the presenceof sulfur-bearing contaminants, of reforming catalysts as describedherein can be determined according to a standard test in which a small,5-100 gram catalyst sample is loaded into a fixed-bed reforming reactorand contacted with a feed blend of 30 mol-% methane, 30 mol-% CO₂, and30 mol-% H₂O that is spiked with 800 mol-ppm of H₂S. In this standardtest, with flowing conditions of 0.7 hr⁻¹ WHSV, a catalyst bedtemperature of 788° C. (1450° F.), and a CO₂-steam reforming reactorpressure of 138 kPa (20 psig), a conversion of the methane of at least85%, and preferably at least 95%, is maintained, at constant catalystbed temperature, for at least 50 hours of operation, and more typicallyfor at least 100 hours of operation, or even for at least 400 hours ofoperation.

The tolerance, or “robustness” of reforming catalysts described hereinis further manifested in a high stability against deactivation in thepresence of other compounds in the gaseous mixture, including highermolecular weight hydrocarbons such as reactive aromatic hydrocarbonsand/or olefinic hydrocarbons that are normally considered prone tocausing reforming catalyst deactivation through coking. For example, thegaseous mixture may comprise aromatic and olefinic hydrocarbons in acombined amount of generally at least about 1 mole-% (e.g., from about 1mol-% to about 25 mol-%), such as at least about 3 mol-% (e.g., fromabout 3 mol-% to about 20 mol-%) or more particularly at least about 5mol-% (e.g., from about 5 mol-% to about 15 mol-%). At such levels ofaromatic and/or olefinic hydrocarbons, reforming catalyst stability maybe exhibited according to the same activity stability test as definedabove with respect to sulfur tolerance, with the exception of the feedblend containing these concentrations of aromatic and/or olefinichydrocarbons as opposed to H₂S. This tolerance of reforming catalysts asdescribed herein with respect to both sulfur and reactive hydrocarbonsallows for the reforming of wide-ranging hydrocarbon-containingfeedstocks, including various fractions (e.g., naphtha and jet fuel)obtained from crude oil refining as described in greater detail below.

More generally, the gaseous mixture, and particularly thehydrocarbon-containing feedstock component of this mixture, maycomprise, in addition to methane, other hydrocarbons such as C₂, C₃,and/or C₄ hydrocarbons (e.g., ethane, propane, propylene, butane, and/orbutenes) that may be present in natural gas and/or other sources ofmethane). Alternatively, reforming catalysts as described herein may beused for dry reforming or CO₂-steam reforming of predominantly, or only,higher molecular weight hydrocarbons, such as in the case of thehydrocarbons in gaseous mixture comprising, or optionally consisting of,any one or more compounds selected from the group consisting of a C₄hydrocarbon, a C₅ hydrocarbon, a C₆ hydrocarbon, a C₇ hydrocarbon, a C₈hydrocarbon, a C₉ hydrocarbon, a C₁₀ hydrocarbon, a C₁₁ hydrocarbon, aC₁₂ hydrocarbon, a C₁₃ hydrocarbon, a C₁₄ hydrocarbon, a C₁₅hydrocarbon, a C₁₆ hydrocarbon, a C₁₇ hydrocarbon, a Cis hydrocarbon,and combinations thereof. For example, the hydrocarbons in the gaseousmixture may comprise, or consist of, C₄-C₈ or C₄-C₆ hydrocarbons, in thecase of dry reforming or CO₂-steam reforming of naphtha boiling-rangehydrocarbons (naphtha reforming). As another example, the hydrocarbonsin the gaseous mixture may comprise, or consist of, C₈-C₁₈ or C₅-C₁₄hydrocarbons, in the case of dry reforming or CO₂-steam reforming of jetfuel boiling-range hydrocarbons (jet fuel reforming). Such naphthaboiling-range hydrocarbons and jet fuel boiling-range fractions arenormally obtained as products from crude oil refining and, as such, canbe a source of sulfur-bearing contaminants in the gaseous mixture. Inrepresentative embodiments, the gaseous mixture may comprise methaneand/or any of the hydrocarbons described herein in a combined amountgenerally from about 5 mol-% to about 85 mol-%, typically from about 10mol-% to about 65 mol-%, and often from about 20 mol-% to about 45mol-%. The gaseous mixture may further comprise CO₂ in an amountgenerally from about 8 mol-% to about 90 mol %, typically from about 15mol-% to about 75 mol-%, and often from about 20 mol-% to about 50mol-%. In the case of CO₂-steam reforming, the gaseous mixture maycomprise H₂O in an amount generally from about 15 mol-% to about 70mol-%, typically from about 20 mol-% to about 60 mol-%, and often fromabout 25 mol-% to about 55 mol-%. The balance of the gaseous mixture mayinclude contaminants such as H₂S and/or other sulfur-bearingcontaminants as described above.

In the case of gaseous mixtures comprising methane and/or lighthydrocarbons (e.g., C₂-C₃ or C₂-C₄ hydrocarbons), the synthesis gasproduct of dry reforming or CO₂-steam reforming may advantageously beused with a favorable molar H₂:CO ratio in the downstream production ofliquid hydrocarbon fuels through Fischer-Tropsch synthesis, as describedabove. The synthesis gas may alternatively be used for other downstreamapplications associated with conventional steam methane reforming (SMR).For example, Tarun (INTERNATIONAL JOURNAL OF GREENHOUSE GAS CONTROL I(2007): 55-61) describes a conventional hydrogen production processinvolving SMR. If dry reforming or CO₂-steam reforming, as describedherein, is applied in hydrogen production, according to embodiments ofthe invention, representative processes may further comprise steps of(i) subjecting the synthesis gas product to one or more water-gas shift(WGS) reaction stages to increase its hydrogen content and/or (ii)separating the effluent of the WGS stage(s), or otherwise separating thesynthesis gas product without intervening WGS stage(s), as the case maybe (e.g., by pressure-swing adsorption (PSA) or membrane separation), toprovide a hydrogen-enriched product stream and a hydrogen-depleted PSAtail gas stream (or simply “PSA tail gas”). The hydrogen-enrichedproduct stream may then be used in a conventional refinery process suchas a hydrotreating process (e.g., hydrodesulfurization, hydrocracking,hydroisomerization, etc.). The hydrogen-depleted PSA tail gas stream maythen be separated to recover hydrogen and/or used as combustion fuel tosatisfy at least some of the heating requirements of the dry reformingor CO₂-steam reforming. In yet further embodiments, the CO— andH₂-containing PSA tail gas may be passed to a biological fermentationstage for the production of fermentation products such as alcohols(e.g., ethanol).

The gaseous effluent from the fermentation stage may then be separatedto recover hydrogen and/or used as combustion fuel as described above.With respect to conventional hydrogen production, the furtherintegration of a biological fermentation stage is described, forexample, in U.S. Pat. Nos. 9,605,286; 9,145,300; US 2013/0210096; and US2014/0028598. As an alternative to integration in a hydrogen productionprocess, dry reforming or CO₂-steam reforming as described herein may beused to provide a synthesis gas product that is used directly in thedownstream production of fermentation products using suitablecarboxydotrophic bacteria (e.g., of the species Clostridiumautoethanogenum or Clostridium ljungdahlii). In either case, i.e., withor without such integration, the microorganisms used for thefermentation may be sulfur tolerant or even require sulfur in the cellculture medium, such that the sulfur tolerance of reforming catalysts asdescribed herein can be particularly advantageous over conventionalreforming catalysts, in terms of compatibility and cost savingsassociated with the elimination of, or at the least reduced requirementsfor, upstream sulfur removal.

Aspects of the invention therefore relate to dry reforming processes andCO₂-steam reforming processes for producing a synthesis gas product(i.e., comprising both H₂ and CO, and optionally other gases such asunconverted CO₂, H₂O, and/or hydrocarbons). In representativeembodiments, a gaseous mixture comprising methane and/or otherhydrocarbon(s) may be provided batchwise, but preferably as a continuousflow, to a reactor of a dry reforming process (i.e., a dry reformingreactor, in the case of the feed or gaseous mixture further comprisingCO₂ but no water) or a CO₂-steam reforming process (i.e., a CO₂-steamreforming reactor, in the case of the feed or gaseous mixture furthercomprising both CO₂ and water), with the general term “reformingreactor” encompassing either case. A synthesis gas product, in turn, maybe withdrawn batchwise (if the gaseous mixture is provided batchwise),but preferably as a continuous flow (if the gaseous mixture is providedas a continuous flow), from the dry reforming reactor or the CO₂-steamreforming reactor, as the case may be.

In addition to H₂, CO, and optionally other gases, water (H₂O) may alsobe present in the synthesis gas product, although at least a portion ofthe water that is present in vapor form may be readily separated bycooling/condensation, for example upstream of a Fischer-Tropschsynthesis reactor (FT reactor) used to convert the synthesis gas productto liquid hydrocarbons. Neither water nor CO₂ in the synthesis gasproduct has an effect on its molar H₂:CO ratio which, as describedabove, is an important parameter in determining the suitability of thesynthesis gas product as a direct feed stream to the FT reactor.

In representative processes, a gaseous mixture comprising methane and/orother light hydrocarbon(s) (e.g., ethane, ethylene, propane, and/orpropylene) and CO₂, as well as optionally H₂O, is contacted with areforming catalyst having activity for carrying out the reforming ofsuch hydrocarbon(s). In particular, such hydrocarbon(s), for example themajority of such hydrocarbons, may be reformed (i) through theiroxidation with some or all of the CO₂ only, according to a dry reformingprocess, or (ii) through their oxidation with both some or all of theCO₂ and some of all of the H₂O (if present), according to a CO₂-steamreforming process.

As described above, aspects of the invention are associated with thediscovery of reforming catalysts for such dry reforming and CO₂-steamreforming processes, exhibiting important advantages, particularly interms of sulfur tolerance and/or a reduced rate of carbon formation(coking), compared to conventional reforming catalysts. Thesecharacteristics, in turn, reduce the rate of catalyst deactivationthrough poisoning and/or coking mechanisms that chemically and/orphysically block active catalyst sites. Further improvements inreforming catalyst stability result at least in part from the highactivity of reforming catalysts described herein, as necessary to lowerthe substantial activation energy barrier associated with the use of CO₂as an oxidant for methane and/or other hydrocarbon(s), as describedabove. This high activity manifests in lower operating (dry reformingreactor or CO₂-steam reforming reactor or dry reforming catalyst bed orCO₂-steam reforming catalyst bed) temperatures, which further contributeto the reduced rate of carbon deposition (coke formation) on thereforming catalyst surface and extended, stable operation. According toparticular embodiments, processes utilizing reforming catalystsdescribed herein can maintain stable operating parameters as describedherein, for example in terms of hydrocarbon conversion (e.g., at leastabout 85% conversion of methane and/or other hydrocarbon(s)) and/ormolar H₂/CO ratio (e.g., from about 1.5:1 to about 2.3:1) of thesynthesis gas product, for at least about 100, at least about 300, oreven at least about 500, hours of continuous or possibly discontinuousoperation. This may be an operating period over which (i) the reformingcatalyst does not undergo regeneration, for example according to areforming process utilizing the catalyst as a fixed bed within thereforming reactor and/or (ii) the temperature of the reforming reactoror respective dry reforming catalyst bed or CO₂-steam reforming catalystbed is not raised beyond a threshold temperature difference from thestart of the time period to the end of the time period, with thisthreshold temperature difference being, for example, 100° C. (180° F.),50° C. (90° F.), 25° C. (45° F.), 10° C. (18° F.), or even 5° C. (9°F.).

Representative reforming catalysts suitable for catalyzing the reactionof methane and/or other hydrocarbon(s) with CO₂ and optionally also withH₂O comprise a noble metal, and possibly two or more noble metals, on asolid support. The solid support preferably comprises a metal oxide,with cerium oxide being of particular interest. Cerium oxide may bepresent in an amount of at least about 80 wt-% and preferably at leastabout 90 wt-%, based on the weight of the solid support (e.g., relativeto the total amount(s) of metal oxide(s) in the solid support). Thesolid support may comprise all or substantially all (e.g., greater thanabout 95 wt-%) cerium oxide. Other metal oxides, such as aluminum oxide,silicon oxide, titanium oxide, zirconium oxide, magnesium oxide,strontium oxide, etc., may also be present in the solid support, incombined amounts representing a minor portion, such as less than about50 wt-%, less than about 30 wt-%, or less than about 10 wt-%, of thesolid support. In other embodiments, the solid support may comprise suchother metal oxides alone or in combination, with a minor portion (e.g.,less than about 50 wt-% or less than about 30 wt-%) of cerium oxide.

Noble metals are understood as referring to a class of metallic elementsthat are resistant to oxidation. In representative embodiments, thenoble metal, for example at least two noble metals, of the reformingcatalyst may be selected from the group consisting of platinum (Pt),rhodium (Rh), ruthenium (Ru), palladium (Pd), silver (Ag), osmium (Os),iridium (Ir), and gold (Au), with the term “consisting of” being usedmerely to denote group members, according to a specific embodiment, fromwhich the noble metal(s) are selected, but not to preclude the additionof other noble metals and/or other metals generally. Accordingly, areforming catalyst comprising a noble metal embraces a catalystcomprising at least two noble metals, as well as a catalyst comprisingat least three noble metals, and likewise a catalyst comprising twonoble metals and a third, non-noble metal such as a promoter metal(e.g., a transition metal). According to preferred embodiments, thenoble metal is present in an amount, or alternatively the at least twonoble metals are each independently present in amounts, from about 0.05wt-% to about 5 wt-%, from about 0.3 wt-% to about 3 wt-%, or from about0.5 wt-% to about 2 wt-%, based on the weight of the catalyst. Forexample, a representative reforming catalyst may comprise the two noblemetals Pt and Rh, and the Pt and Rh may independently be present in anamount within any of these ranges (e.g., from about 0.05 wt-% to about 5wt-%). That is, either the Pt may be present in such an amount, the Rhmay be present in such an amount, or both Pt and Rh may be present insuch amounts.

In representative embodiments, the at least two noble metals (e.g., Ptand Rh) may be substantially the only noble metals present in thereforming catalyst, such that, for example, any other noble metal(s)is/are present in an amount or a combined amount of less than about 0.1wt-%, or less than about 0.05 wt-%, based on the weight of the reformingcatalyst. In further representative embodiments, that at least two noblemetals (e.g., Pt and Rh) are substantially the only metals present inthe reforming catalyst, with the exception of metals present in thesolid support (e.g., such as cerium being present in the solid supportas cerium oxide). For example, any other metal(s), besides at least twonoble metals and metals of the solid support, may be present in anamount or a combined amount of less than about 0.1 wt-%, or less thanabout 0.05 wt-%, based on the weight of the reforming catalyst. Anymetals present in the catalyst, including noble metal(s), may have ametal particle size in the range generally from about 0.3 nanometers(nm) to about 20 nm, typically from about 0.5 nm to about 10 nm, andoften from about 1 nm to about 5 nm.

The noble metal(s) may be incorporated in the solid support according toknown techniques for catalyst preparation, including sublimation,impregnation, or dry mixing. In the case of impregnation, which is apreferred technique, an impregnation solution of a soluble compound ofone or more of the noble metals in a polar (aqueous) or non-polar (e.g.,organic) solvent may be contacted with the solid support, preferablyunder an inert atmosphere. For example, this contacting may be carriedout, preferably with stirring, in a surrounding atmosphere of nitrogen,argon, and/or helium, or otherwise in a non-inert atmosphere, such asair. The solvent may then be evaporated from the solid support, forexample using heating, flowing gas, and/or vacuum conditions, leavingthe dried, noble metal-impregnated support. The noble metal(s) may beimpregnated in the solid support, such as in the case of two noblemetals being impregnated simultaneously with both being dissolved in thesame impregnation solution, or otherwise being impregnated separatelyusing different impregnation solutions and contacting steps. In anyevent, the noble metal-impregnated support may be subjected to furtherpreparation steps, such as washing with the solvent to remove excessnoble metal(s) and impurities, further drying, calcination, etc. toprovide the reforming catalyst.

The solid support itself may be prepared according to known methods,such as extrusion to form cylindrical particles (extrudates) or oildropping or spray drying to form spherical particles. Regardless of thespecific shape of the solid support and resulting catalyst particles,the amounts of noble metal(s) being present in the reforming catalyst,as described above, refer to the weight of such noble metal(s), onaverage, in a given catalyst particle (e.g., of any shape such ascylindrical or spherical), independent of the particular distribution ofthe noble metals within the particle. In this regard, it can beappreciated that different preparation methods can provide differentdistributions, such as deposition of the noble metal(s) primarily on ornear the surface of the solid support or uniform distribution of thenoble metal(s) throughout the solid support. In general, weightpercentages described herein, being based on the weight of the solidsupport or otherwise based on the weight of reforming catalyst, canrefer to weight percentages in a single catalyst particle but moretypically refer to average weight percentages over a large number ofcatalyst particles, such as the number in a reforming reactor that forma catalyst bed as used in processes described herein.

Simplified illustrations of dry reforming processes and optionallyCO₂-steam reforming processes 10 are depicted in FIGS. 1A and 1B. Ineither of these embodiments, gaseous mixture 4 comprising one or morehydrocarbons (e.g., methane) and CO₂, may reside within reformingreactor 5 in the form of a vessel that is used to contain a bed ofreforming catalyst 6, as described above, under reforming conditions atwhich gaseous mixture 4 and reforming catalyst 6 are contacted.According to the embodiment illustrated in FIG. 1A, gaseous mixture 4may be provided within reforming reactor 5 from hydrocarbon-containingfeedstock 1 alone. For example, a representative hydrocarbon-containingfeedstock is a methane-containing feedstock that is obtained frombiomass gasification or pyrolysis, including hydrogasification orhydropyrolysis, and may further comprise CO₂ and H₂O. Such ahydrocarbon-containing feedstock may thereby itself provide gaseousmixture 4 for a CO₂-steam reforming process, in which both CO₂ and H₂Oreact as oxidants of methane. In other embodiments, gaseous mixture 4may be obtained from combining hydrocarbon-containing feedstock 1 withoptional CO₂-containing oxidant 2, if, for example,hydrocarbon-containing feedstock 1 contains little CO₂ such as in thecase of liquid hydrocarbons including naphtha boiling-range hydrocarbonsand/or jet fuel boiling-range hydrocarbons, or otherwise in the case ofsome types of natural gas.

As another option, H₂O-containing oxidant 3 (e.g., as steam) may also becombined to form gaseous mixture 4, comprising methane and both CO₂ andH₂O oxidants for a CO₂-steam reforming processes. Again, however, H₂Omay also be present in sufficient quantity in hydrocarbon-containingfeedstock 1 and/or CO₂-containing oxidant 2, such that separateH₂O-containing oxidant 3 may not be necessary. As shown by dashed,double-headed arrows between hydrocarbon-containing feedstock 1,CO₂-containing oxidant 2, and H₂O-containing oxidant 3, it is clear thatany of these may be combined prior to (e.g., upstream of) reformingreactor 5. According to a specific embodiment, FIG. 1B illustrateshydrocarbon-containing feedstock 1 being combined with optionalCO₂-containing oxidant 2 and optional H₂O-containing oxidant 3 toprovide gaseous mixture 4 both prior to (e.g., upstream of) reformingreactor 5, as well as within this reactor.

As described above, in embodiments in which gaseous mixture 4 comprisesone or more hydrocarbons such as methane and CO₂, but not H₂O, theprocess may be considered a “dry reforming” process, whereas inembodiments in which gaseous mixture 4 comprises hydrocarbon(s) and CO₂,and further comprises H₂O acting, in combination with the CO₂, asoxidants of the hydrocarbon(s) (e.g., such that at least respectiveoxidant portions of the CO₂ and H₂O oxidize respective reactant portionsof the hydrocarbon(s)), the process may be considered a “CO₂-steamreforming process.” Reforming catalysts as described herein provideadvantageous results in both dry reforming and CO₂-steam reforming, interms of both activity and stability, as described above. Underreforming conditions provided in reforming reactor 5, gaseous mixture 4is converted to synthesis gas product 7, which may, relative to gaseousmixture 4, be enriched in (i.e., have a higher concentration of)hydrogen and CO, and/or be depleted in (i.e., have a lower concentrationof) CO₂, H₂O, methane, and/or other hydrocarbon(s) initially present ingaseous mixture 4.

An important methane-containing feedstock is natural gas, andparticularly stranded natural gas, which, using known processes, is noteasily converted to a synthesis gas product in an economical manner.Natural gas comprising a relatively high concentration of CO₂, forexample at least about 10 mol-% or even at least about 25 mol-%,represents an attractive methane-containing feedstock, since processesas described herein do not require the removal of CO₂ (e.g., byscrubbing with an amine solution), in contrast to conventional steamreforming, and in fact utilize CO₂ as a reactant. Othermethane-containing feedstocks may comprise methane obtained from coal orbiomass (e.g., lignocellulose or char) gasification, from a biomassdigester, or as an effluent from a renewable hydrocarbon fuel (biofuel)production processes (e.g., a pyrolysis process, such as ahydropyrolysis processes, or a fatty acid/triglyceride hydroconversionprocesses). Further methane-containing feedstocks may comprise methaneobtained from a well head or an effluent of an industrial processincluding a petroleum refining process (as a refinery off gas), anelectric power production process, a steel manufacturing process or anon-ferrous manufacturing process, a chemical (e.g., methanol)production process, or a coke manufacturing process. Generally, anyprocess gas known to contain a hydrocarbon (e.g., a C₁-C₃ hydrocarbon)and CO₂ may provide all or a portion of the gaseous mixture as describedherein, or at least all or a portion of the methane-containing feedstockas a component of this mixture. If the methane-containing feedstockcomprises methane obtained from a renewable resource (e.g., biomass),for example methane from a process stream obtained by hydropyrolysis asdescribed in U.S. Pat. No. 8,915,981 assigned to Gas TechnologyInstitute, then processes described herein may be used to producerenewable synthesis gas products (i.e., comprising renewable CO) that,in turn, can be further processed to provide renewablehydrocarbon-containing fuels, fuel blending components, and/orchemicals. Accordingly, the methane-containing feedstock may thereforecomprise methane from a non-renewable source (e.g., natural gas) and/ormethane from a renewable source (e.g., biomass), with the latter sourceimparting an overall reduction in the carbon footprint associated withthe synthesis gas product and downstream products. As further describedherein, natural gas and/or other methane-containing feedstocks, may be,but need not be, pretreated to remove H₂S and other sulfur-bearingcontaminants, prior to dry reforming or CO₂-steam reforming.

Like the methane-containing feedstock (or hydrocarbon-containingfeedstock generally), and particularly in view of the sulfur toleranceof reforming catalysts as described herein, other components of thegaseous mixture, including the CO₂-containing oxidant and/orH₂O-containing oxidant, may be obtained from a wide variety of sources.Advantageously, such sources include waste gases that are regarded ashaving little or no economic value, and that may additionally contributeto atmospheric CO₂ levels. For example, the CO₂-containing oxidant maycomprise an industrial process waste gas that is obtained from a steelmanufacturing process or a non-ferrous product manufacturing process.Other processes from which all or a portion of the CO₂-containingoxidant may be obtained include petroleum refining processes, renewablehydrocarbon fuel (biofuel) production processes (e.g., a pyrolysisprocess, such as a hydropyrolysis processes, or a fattyacid/triglyceride hydroconversion processes), coal and biomassgasification processes, electric power production processes, carbonblack production processes, ammonia production processes, methanolproduction processes, and coke manufacturing processes.

As described above, the methane-containing feedstock (orhydrocarbon-containing feedstock generally) may itself provide thegaseous mixture for a dry reforming process or a CO₂-steam reformingprocess, i.e., without the addition of a separate CO₂-containing oxidantand/or a separate H₂O-containing oxidant, if sufficient CO₂ and/or H₂Oare already present in this mixture. Alternatively, themethane-containing feedstock (or hydrocarbon-containing feedstockgenerally), may be combined with only one of a CO₂-containing oxidant orH₂O-containing oxidant to provide a suitable gaseous mixture. Forexample, steam (as the H₂O-containing oxidant) may be combined with amethane-containing feedstock further comprising CO₂, to provide agaseous mixture suitable for a CO₂-steam reforming process.

A representative methane-containing feedstock further comprising CO₂ inan amount particularly suitable for providing the gaseous mixture for aCO₂-steam reforming process described herein is a hydropyrolysis gaseousmixture obtained from biomass hydropyrolysis and having (i) a methaneconcentration of generally about 3 mol-% to about 45 mol-% (e.g., about5 mol-% to about 25 mol-% or about 7 mol-% to about 15 mol-%), (ii)ethane and propane concentrations each of generally about 1 mol-% toabout 35 mol-% (e.g., about 2 mol-% to about 25 mol-% each or about 3mol-% to about 15 mol-% each), and (iii) a CO₂ concentration ofgenerally about 10 mol-% to about 75 mol-% (e.g., about 12 mol-% toabout 55 mol-% or about 15 mol-% to about 35 mol-%). The substantialbalance of the hydropyrolysis gaseous mixture may be water vapor.However, depending on the actual amount of water vapor, anH₂O-containing oxidant may optionally be combined with thehydropyrolysis gaseous mixture to provide the gaseous mixture to aCO₂-steam reforming reactor with a desired molar H₂O:CO₂ ratio. In thiscase, the H₂O-containing oxidant may be readily available as a condensedaqueous phase that is separated from the substantially fullydeoxygenated hydrocarbon liquid generated from the hydropyrolysis ofbiomass (e.g., a hydrocarbon-containing liquid having a total oxygencontent of less than about 2% by weight, or less than about 1% byweight).

Another example of a representative methane-containing feedstock furthercomprising CO₂, in an amount particularly suitable for providing thegaseous mixture for a CO₂-steam reforming process described herein, isnatural gas comprising CO₂ at a concentration of generally about 3 mol-%to about 35 mol-% (e.g., about 5 mol-% to about 30 mol-% or about 10mol-% to about 25 mol-%) and methane at a concentration of generallyabout 65 mol-% to about 98 mol-% (e.g., about 70 mol-% to about 95 mol-%or about 75 mol-% to about 90 mol-%). Other hydrocarbons (e.g., ethaneand propane), as well as nitrogen, may be present in minor amounts. AnH₂O-containing oxidant may optionally be combined with thismethane-containing feedstock to provide the gaseous mixture to aCO₂-steam reforming reactor with a desired molar H₂O:CO₂ ratio.

Another example of a representative methane-containing feedstock furthercomprising CO₂, in an amount particularly suitable for providing thegaseous mixture for a CO₂-steam reforming process described herein, isbiogas obtained from the bacterial digestion of organic waste, such asfrom anaerobic digestion processes and from landfills. Biogas containsmethane at a concentration of generally about 35 mol-% to about 90 mol-%(e.g., about 40 mol-% to about 80 mol-% or about 50 mol-% to about 75mol-%) and CO₂ at a concentration of generally about 10 mol-% to about60 mol-% (e.g., about 15 mol-% to about 55 mol-% or about 25 mol-% toabout 50 mol-%). The gases N₂, H₂, H₂S, and O₂ may be present in minoramounts (e.g., in a combined amount of less than 20 mol-%, or less than10 mol-%). An H₂O-containing oxidant may optionally be combined withthis methane-containing feedstock to provide the gaseous mixture to aCO₂-steam reforming reactor with a desired molar H₂O:CO₂ ratio.

Another example of a representative methane-containing feedstock furthercomprising CO₂, in an amount particularly suitable for providing thegaseous mixture for a CO₂-steam reforming process described herein, is ahydrogen-depleted PSA tail gas, for example obtained from a hydrogenproduction processes involving SMR, as described above. This stream mayhave (i) a methane concentration of generally about 5 mol-% to about 45mol-% (e.g., about 10 mol-% to about 35 mol-% or about 15 mol-% to about25 mol-%), (ii) a CO₂ concentration of generally about 20 mol-% to about75 mol-% (e.g., about 25 mol-% to about 70 mol-% or about 35 mol-% toabout 60 mol-%), and (iii) an H₂ concentration of generally about 10mol-% to about 45 mol-% (e.g., about 15 mol-% to about 40 mol-% or about20 mol-% to about 35 mol-%). The balance of this stream may comprisepredominantly water vapor and/or CO. An H₂O-containing oxidant mayoptionally be combined with this methane-containing feedstock to providethe gaseous mixture to a CO₂-steam reforming reactor with a desiredmolar H₂O:CO₂ ratio.

Another example of a representative methane-containing feedstock furthercomprising CO₂ in an amount particularly suitable for providing thegaseous mixture for a CO₂-steam reforming process described herein is agaseous effluent from a bacterial fermentation that is integrated with ahydrogen production process, as described above. This stream may have(i) a methane concentration of generally about 5 mol-% to about 55 mol-%(e.g., about 5 mol-% to about 45 mol-% or about 10 mol-% to about 40mol-%), (ii) a CO₂ concentration of generally about 5 mol-% to about 75mol-% (e.g., about 5 mol-% to about 60 mol-% or about 10 mol-% to about50 mol-%), and (iii) an H₂ concentration of generally about 5 mol-% toabout 40 mol-% (e.g., about 5 mol-% to about 30 mol-% or about 10 mol-%to about 25 mol-%). The balance of this stream may comprisepredominantly water vapor and/or CO. An H₂O-containing oxidant mayoptionally be combined with this methane-containing feedstock to providethe gaseous mixture to a CO₂-steam reforming reactor with a desiredmolar H₂O:CO₂ ratio.

In representative embodiments, according to FIGS. 1A and 1B, gaseousmixture 4 comprising a hydrocarbon and CO₂ may be contacted withreforming catalyst 6 in a batchwise or discontinuous operation, butpreferably the dry reforming or CO₂-steam reforming process is performedcontinuously with flowing streams of the gaseous mixture 4 or componentsthereof (e.g., hydrocarbon-containing feedstock 1, CO₂-containingoxidant 2, and/or H₂O-containing oxidant 3 as described herein), toimprove process efficiency. For example, contacting may be performed bycontinuously flowing the gaseous mixture 4 (e.g., as a combinedreforming reactor feed stream of any of these components in combination)through the reforming reactor 5 and reforming catalyst 6 under reformingconditions (e.g., conditions within a reforming reactor vessel andwithin a bed of the reforming catalyst that is contained in the vessel)that include a suitable flow rate. In particular embodiments, thereforming conditions may include a weight hourly space velocity (WHSV)generally from about 0.05 hr⁻¹ to about 10 hr⁻¹, typically from about0.1 hr⁻¹ to about 4.0 hr⁻¹, and often from about 0.3 hr⁻¹ to about 2.5hr⁻¹. As is understood in the art, the WHSV is the weight flow of atotal feed (e.g. the gaseous mixture) to a reactor, divided by theweight of the catalyst in the reactor and represents the equivalentcatalyst bed weights of the feed stream processed every hour. The WHSVis related to the inverse of the reactor residence time. The reformingcatalyst 6 may be contained within reforming reactor 5 in the form of afixed bed, but other catalyst systems are also possible, such as movingbed and fluidized bed systems that may be beneficial in processes usingcontinuous catalyst regeneration.

Other reforming conditions, which are useful for either dry reforming orCO₂-steam reforming, include a temperature generally from about 649° C.(1200° F.) to about 816° C. (1500° F.). Processes described herein, byvirtue of the high activity of the reforming catalyst in terms ofreducing the activation energy barrier required for the use of CO₂ as anoxidant, can effectively oxidize methane and/or other hydrocarbons atsignificantly lower temperatures, compared to a representativeconventional temperature of 950° C. (1742° F.) that is used for dryreforming or steam reforming. For example, in representativeembodiments, the reforming conditions can include a temperature in arange from about 677° C. (1250° F.) to about 788° C. (1450° F.), or fromabout 704° C. (1300° F.) to about 760° C. (1400° F.). As describedabove, the presence of H₂S and/or other sulfur-bearing contaminants insignificant amounts (e.g., 100-1000 mol-ppm) may warrant increasedtemperatures, for example in a range from about 732° C. (1350° F.) toabout 843° C. (1550° F.), or from about 760° C. (1400° F.) to about 816°C. (1500° F.), to maintain desired conversion levels (e.g., greater thanabout 85%). Yet other reforming conditions can include an above-ambientpressure, i.e., a pressure above a gauge pressure of 0 kPa (0 psig),corresponding to an absolute pressure of 101 kPa (14.7 psia). Becausethe reforming reactions make a greater number of moles of product versusmoles of reactant, equilibrium is favored at relatively low pressures.Therefore, reforming conditions can include a gauge pressure generallyfrom about 0 kPa (0 psig) to about 517 kPa (75 psig), typically fromabout 0 kPa (0 psig) to about 345 kPa (50 psig), and often from about103 kPa (15 psig) to about 207 kPa (30 psig).

Advantageously, within any of the above temperature ranges, the highactivity of the reforming catalyst can achieve a conversion of methaneand/or other hydrocarbon(s) (e.g., a conversion of methane, a conversionof combined C₁-C₃ hydrocarbons, a conversion of combined C₁-C₄hydrocarbons, a conversion of naphtha boiling-range hydrocarbons, aconversion of jet fuel boiling-range hydrocarbons, etc.) of at leastabout 80% (e.g., from about 80% to about 99%), at least about 85% (e.g.,from about 85% to about 97%), or at least about 90% (e.g., from about90% to about 99%), for example by adjusting the particular reformingreactor temperature or reforming catalyst bed temperature and/or otherreforming conditions (e.g., WHSV and/or pressure) as would beappreciated by those having skill in the art, with knowledge gained fromthe present disclosure. Advantageously, reforming catalysts as describedherein are sufficiently active to achieve a significant hydrocarbon(e.g., methane) conversion, such as at least about 85%, in a stablemanner at a temperature of at most about 732° C. (1350T), or even atmost about 704° C. (1300° F.). With respect to the oxidant reactants, arepresentative conversion of CO₂ is at least about 50% (e.g., from about50% to about 75%), and a representative conversion of H₂O is at leastabout 70% (e.g., from about 70% to about 90%), at the conversion levelsdescribed herein with respect to hydrocarbon(s). As is understood in theart, conversion of any particular compound (e.g., methane) orcombination of compounds (e.g., C₁-C₄ hydrocarbons or C₁-C₃hydrocarbons) can be calculated on the basis of:

100*(X _(feed) −X _(prod))/X _(feed),

wherein X_(feed) is the total amount (e.g., total weight or total moles)of the compound(s) X in the gaseous mixture (e.g., combined reactorfeed) provided to a reactor and X_(prod) is the total amount of thecompound(s) X in the synthesis gas product removed from the reactor. Inthe case of continuous processes, these total amounts may be moreconveniently expressed in terms of flow rates, or total amounts per unittime (e.g., total weight/hr or total moles/hr). Other performancecriteria that can be achieved using reforming catalysts and reformingconditions as described herein include a high hydrogen yield, or portionof the total hydrogen in the methane and/or other hydrogen-containingcompounds (e.g., total hydrogen in the hydrocarbons such as C₂-C₄hydrocarbons or C₂-C₃ hydrocarbons), in the gaseous mixture provided tothe reactor, which is converted to H₂ in the synthesis gas productremoved from the reactor. In representative embodiments, the hydrogenyield is at least about 70% (e.g., from about 70% to about 85%). Asdescribed above with respect to conversion, amounts provided to andremoved from the reactor may be expressed in terms of flow rates.

As described above, further advantages associated with reformingprocesses, and particularly CO₂-steam reforming processes, as describedherein, include favorable molar H₂/CO ratios, as well as the ability toadjust these ratios, in the synthesis gas product. This has especiallyimportant implications for downstream processing via Fischer-Tropsch forthe production of liquid hydrocarbons. The exact composition of thesynthesis gas product depends on the composition of the feed (e.g.,combined reforming reactor feed) or gaseous mixture, the reformingcatalyst, and the reforming conditions.

In representative embodiments, the synthesis gas product, particularlyin the case of a CO₂-steam reforming process, advantageously has a molarH₂:CO ratio that is near 2:1, for example generally in a range fromabout 1.5:1 to about 2.3:1, and typically from about 1.8:1 to about2.2:1. The combined concentration of 1H₂ and CO in this product isgenerally at least about 35 mol-% (or vol-%) (e.g., from about 35 mol-%to about 85 mol-%), typically at least about 50 mol-% (e.g., from about50 mol-% to about 80 mol-%), and often at least about 60 mol-% (e.g.,from about 60 mol-% to about 75 mol-%). As described above, the balanceof the synthesis gas product may be substantially or all CO₂ and water,depending on the particular dry reforming or CO₂-steam reformingprocess, including the conditions of such process (e.g., conditionswithin the reforming reactor such as temperature, pressure, weighthourly space velocity, and reforming catalyst formulation) and the feedor gaseous mixture being reacted. In representative embodiments, CO₂ ispresent in the synthesis gas product in a concentration of generallyless than about 45 mol-% (e.g., from about 5 mol-% to about 45 mol-%)and typically less than about 35 mol-% (e.g., from about 10 mol-% toabout 35 mol-%). Water may be present in a concentration of generallyless than about 20 mol-% (e.g., from about 1 mol-% to about 25 mol-%)and typically less than about 15 mol-% (e.g., from about 5 mol-% toabout 15 mol-%). Minor amounts of unconverted hydrocarbons may also bepresent in the synthesis gas product. For example, a combined amount ofC₁-C₄ hydrocarbons (e.g., a combined amount of methane, ethane, propane,and butane), which may possibly include only C₁-C₃ hydrocarbons, may bepresent in a concentration of less than about 5 mol-% and typically lessthan about 2 mol-%.

Integrated Processes Including Conversion Steps to Produce LiquidHydrocarbons

Further representative processes use dry reforming or CO₂-steamreforming, as described herein, with additional process steps, such asconverting H₂ and CO in the synthesis gas product in an FT synthesisstage, in order to provide a Fischer-Tropsch product (e.g., effluentfrom an FT reactor as described above) comprising hydrocarbons,including C₄ ⁺ hydrocarbons representative of those present in liquidfuels such as gasoline, jet fuel, and/or diesel fuel. For example, aparticular integrated process for producing C₄ ⁺ hydrocarbons maycomprise, in a reforming reactor of a reforming stage, convertingmethane and CO₂ in a gaseous mixture, such as any of the gaseousmixtures described herein, including gaseous mixtures that may compriseany methane-containing feedstock or other component of such gaseousmixture as described above, to produce a synthesis gas product asdescribed above. This converting step may more particularly comprisecontacting the gaseous mixture with a reforming catalyst, such as any ofthe reforming catalysts described herein, in a reforming reactor of areforming stage to produce the synthesis gas product. The integratedprocess may further comprise, in an FT reactor of an FT synthesis stagedownstream of the reforming stage, converting H₂ and CO in the synthesisgas product to hydrocarbons, including C₄ ⁺ hydrocarbons (i.e., at leastsome hydrocarbons having four or more carbon atoms) that are provided inan FT product. As an optional step, and particularly in the case of theC₄ ⁺ hydrocarbons in the FT product including a wax fraction comprisingnormal C₂₀ ⁺ hydrocarbons (i.e., at least some normal or straight-chainhydrocarbons having 20 or more carbon atoms that are consequently solidat room temperature), the integrated process may further comprise, in afinishing reactor of a finishing stage downstream of the FT synthesisstage, converting at least a portion of the normal C₂₀ ⁺ hydrocarbons tonormal or branched C₄-C₁₉ hydrocarbons (i.e., to normal or branchedhydrocarbons, at least some of which have 4 to 19 carbon atoms) that areprovided in a hydroisomerization/hydrocracking product.

The term “stage” as used in “reforming stage,” “FT synthesis stage,” and“finishing stage,” refers to reactor(s) used to carry out the reactionsassociated with these stages as described herein, as well as thecatalyst(s) and conventional auxiliary equipment (e.g., sensors, valves,gauges, control systems, etc.) associated with the reactor(s). In someembodiments, and preferably, only a single reactor is needed for a givenstage, i.e., a single reforming reactor, a single FT reactor, and/or asingle finishing reactor. However, reactions associated with a givenstage may also be carried out in more than one reactor, for example tworeactors operating in parallel or in series.

Additional details and advantages, in representative integratedprocesses, of the reforming stage, FT synthesis stage, and optionalfinishing stage are provided below, with the understanding thatintegrated processes according to the present disclosure include thosehaving any one of these additional details and/or advantages, orotherwise any combination of such details and/or advantages.

Reforming Stage

The reforming stage includes at least one, and typically only one,reforming reactor as described above, which may be a dry reformingreactor or a CO₂-steam reforming reactor, with the latter termindicating the presence of steam in the gaseous mixture. Gaseousmixtures that are converted in this stage are as described above, aswell as representative reforming catalysts and their properties (e.g.,activity, stability, tolerance to sulfur and higher molecular weighthydrocarbons, etc.), reforming conditions suitable for use in at leastone reforming reactor, and performance criteria (conversion levels andproduct yields).

As described above, the gaseous mixture may be pretreated, upstream ofthe reforming reactor(s), to reduce the concentration of H₂S and/orother sulfur-bearing contaminants, for example by contacting the gaseousmixture or any component thereof (e.g., the hydrocarbon-containingfeedstock) with a suitable bed of sorbent or a liquid wash.Alternatively, a post-treatment (downstream of the reforming stage) ofthe synthesis gas product or possibly of the FT feed (e.g., followingcondensing of water from a cooled synthesis gas product to provide theFT feed) may be performed, for example in this manner, to reduce theconcentration of H₂S and/or other sulfur-bearing contaminants. Theoption to perform a step of removing sulfur-bearing contaminants eitherupstream or downstream of the reforming reactor(s) arises from thesulfur tolerance of reforming catalysts as described above, such thatthe protection of the reforming catalyst from sulfur poisoning may notbe necessary, although protection of the FT catalyst may be necessary.Advantageously, if the concentration of H₂S and/or other sulfur-bearingcontaminants is reduced upstream of the reforming reactor(s) (e.g., anH₂S removal pretreatment is performed on the gaseous mixture), suchpretreatment may be less rigorous and/or involve less gas removal,compared to conventional acid gas removal (e.g., using amine scrubbing)in which CO₂ would also normally be removed. The ability of reformingcatalysts described herein to tolerate CO₂, and in fact utilize this gasas a reactant, can therefore allow for a reduction, or even theelimination, of conventional pretreatment steps. For example, a gaseousmixture comprising natural gas and having a high concentration of CO₂(e.g., greater than 25 mol-% or greater than 30 mol-%), which may be dueto the particular source of the natural gas, may be provided to thereforming reactor(s) without any pretreatment, or possibly with only apretreatment for the removal of dust particles, such as by filtration.

As described above, the reforming stage produces a synthesis gas productcomprising H₂ and CO, by virtue of reacting a hydrocarbon by dryreforming or by CO₂-steam reforming. As further described above, in viewof the favorable ranges of molar H₂:CO ratios (e.g., encompassing 2:1 inthe case of CO₂-steam reforming) of the synthesis gas product that maybe obtained, advantageously some or all of the synthesis gas product maybe used directly in the FT synthesis stage, without any interveningoperation that would impact the molar H₂:CO ratio (e.g., by theaddition, removal, or conversion of components that would alter thisratio, such as by the use of a separate water-gas shift reaction orreverse water-gas shift reaction). Further advantages associated withthe composition of the synthesis gas product are described according toembodiments presented herein and relating to the downstream processingof this product.

FT Synthesis Stage

In the FT reactor(s) or overall FT synthesis stage, at least a portionof the H₂ and CO in the synthesis gas product are converted tohydrocarbons, according to the Fischer-Tropsch (FT) synthesis reactiongiven above. In particular, an FT feed comprising some or all of thesynthesis gas product, optionally following one or more interveningoperations such as cooling, heating, pressurizing, depressurizing,separation of one or more components (e.g., removal of condensed water),addition of one or more components (e.g., addition of H₂ and/or CO toadjust the molar H₂:CO ratio of the FT feed relative to that of thesynthesis gas product), and/or reaction of one or more components (e.g.,reaction of H₂ and/or CO using a separate water-gas shift reaction orreverse water-gas shift reaction), is provided to the FT reactor(s) ofthe FT synthesis stage. In view of the temperatures and pressurestypically used in the FT reactor(s) of the FT synthesis stage relativeto those used in the reforming reactor(s) of the reforming stage, thesynthesis gas product may be cooled, separated from condensed water, andpressurized. In some embodiments, these may be the only interveningoperations to which the synthesis gas product is subjected, to providethe FT feed. In other embodiments, cooling and pressurizing may be theonly intervening operations. In yet other embodiments, interveningoperations that may be omitted include drying of the synthesis gasproduct to remove vapor phase H₂O (which is therefore different fromcondensing liquid phase H₂O and can include, e.g., using a sorbentselective for water vapor, such as 5A molecular sieve) and/or CO₂removal according to conventional acid gas treating steps (e.g., aminescrubbing). However, according to some embodiments, CO₂ removal may beperformed downstream of the reforming stage but upstream of the FTsynthesis stage (e.g., as an intervening operation), in lieu ofperforming this CO₂ removal upstream of the reforming stage, as isconventionally practiced. Preferably, prior to the FT reactor(s), waterproduced in the reforming reactor is condensed from the synthesis gasproduct, and/or also preferably the molar H₂:CO ratio of the synthesisgas product is not adjusted. The use of no intervening operationsbetween the reforming stage and the FT synthesis stage, limitedintervening operations, and/or the omission or certain interveningoperations, results in advantages associated with the overallsimplification of the integrated process.

Conditions in the FT reactor(s) are suitable for the conversion of H₂and CO to hydrocarbons, including C₄ ⁺ hydrocarbons that are useful asliquid fuels or blending components of liquid fuels. In representativeembodiments, FT reaction conditions (suitable for use in at least one FTreactor) can include a temperature in a range from about 121° C. (250°F.) to about 288° C. (550° F.), or from about 193° C. (380° F.) to about260° C. (500° F.). Other FT reaction conditions can include a gaugepressure from about 689 kPa (100 psig) to about 3.44 MPa (500 psig), orfrom about 1.38 MPa (200 psig) to about 2.76 MPa (400 psig). Oneadvantage over the use of an FT synthesis stage downstream of thereforming stage, relative to the downstream production of methanoland/or DME as described above, is the significantly reduced pressure(e.g., generally below about 3.44 MPa (500 psig) or typically belowabout 3.10 MPa (450 psig)) compared to these downstream processingalternatives.

In the FT reactor(s), the FT feed may be contacted with a suitable FTcatalyst (e.g., bed of FT catalyst particles disposed within the FTreactor) under FT reaction conditions, which may include thetemperatures and/or pressures as described above. Representative FTcatalysts comprise one or more transition metals selected from cobalt(Co), iron (Fe), ruthenium (Ru), and nickel (Ni). A preferred FTcatalyst comprises at least about 10 wt-% of the transition metal(s),and typically at least about 15 wt-% of the transition metal(s), on asolid support. The phrase “on a solid support” is intended to encompasscatalysts in which the active metal(s) is/are on the support surfaceand/or within a porous internal structure of the support. Representativesolid supports comprise one or more metal oxides, selected from thegroup consisting of aluminum oxide, silicon oxide, titanium oxide,zirconium oxide, magnesium oxide, strontium oxide, etc. The solidsupport may comprise all or substantially all (e.g., greater than about95 wt-%) of the one or more of such metal oxides. Preferred FT catalystscomprise the transition metal cobalt (Co) in the above amounts (e.g., atleast about 10 wt-%) on a support comprising aluminum oxide (alumina).

The FT catalysts and FT reaction conditions described herein aregenerally suitable for achieving a conversion of H₂ and/or CO (H₂conversion or CO conversion) of at least about 20% (e.g., from about 20%to about 99% or from about 20% to about 75%), at least about 30% (e.g.,from about 30% to about 95% or from about 30% to about 65%), or at leastabout 50% (e.g., from about 50% to about 90% or from about 50% to about85%). These FT conversion levels may be based on H₂ conversion or COconversion, depending on which reactant is stoichiometrically limited inthe FT feed, considering the FT synthesis reaction chemistry, and theseFT conversion levels may be calculated as described above. Preferably,these FT conversion levels are based on CO conversion. These FTconversion levels may be based on “per-pass” conversion, achieved in asingle pass through the FT synthesis stage (e.g., an FT reactor of thisstage), or otherwise based on overall conversion, achieved by returninga recycle portion of the FT product back to the FT synthesis stage(e.g., an FT reactor of this stage), as described in greater detailbelow.

A desired H₂ conversion and/or CO conversion in the FT reactor(s) may beachieved by adjusting the FT reaction conditions described above (e.g.,FT reaction temperature and/or pressure), and/or adjusting the weighthourly space velocity (WHSV), as defined above. The FT reactionconditions may include a weight hourly space velocity (WHSV) generallyfrom about 0.01 hr⁻¹ to about 10 hr⁻¹, typically from about 0.05 hr⁻¹ toabout 5 hr⁻¹, and often from about 0.3 hr⁻¹ to about 2.5 hr⁻¹. Theconversion level (e.g., CO conversion) may be increased, for example, byincreasing pressure and decreasing WHSV, both of which have the effectof increasing reactant concentrations and reactor residence times. Anexample of the effect of pressure on the level of CO conversion achievedin a Fischer-Tropsch (FT) reactor, containing an FT catalyst asdescribed herein and also while operating with other FT reactionconditions constant and within ranges as described above, is depicted inFIG. 2 . The FT reaction conditions may optionally include returning arecycle portion of the FT product, exiting the FT reactor, back to theFT feed for combining with the FT feed, or otherwise back to the FTreactor itself. Recycle operation allows for operation at relatively low“per-pass” conversion through the FT reactor, while achieving a highoverall conversion due to the recycle. In some embodiments, this lowper-pass conversion may advantageously limit the quantity of highmolecular weight hydrocarbons (e.g., normal C₂₀ ⁺ hydrocarbons) that canbe produced as part of the hydrocarbon product distribution obtainedfrom the FT synthesis reaction.

Preferably, however, the FT reaction conditions include little or evenno FT product recycle. For example, the FT reaction conditions mayinclude a weight ratio of recycled FT product to FT feed (i.e., a“recycle ratio”), with this recycled FT product and FT feed togetherproviding a combined feed to the FT reactor, of generally less thanabout 1:1, typically less than about 0.5:1, and often less than about0.1:1. For example, the recycle ratio may be 0, meaning that no FTproduct recycle is used, such that the per-pass conversion is equal tothe overall conversion. With such low recycle ratios, a relatively highper-pass H₂ conversion or CO conversion, such as at least about 50%(e.g., from about 50% to about 95%), at least about 70% (e.g., fromabout 70% to about 92%), or at least about 80% (e.g., from about 80% toabout 90%), is desirable in view of process efficiency and economics. Asthe per-pass conversion level is increased, the distribution ofhydrocarbons in the FT product is shifted to those having increasednumbers of carbon atoms. This is advantageous in terms of the reductionin yield of light, C₁-C₃ hydrocarbons, having less value than thedesired C₄ ⁺ liquid hydrocarbons. In some embodiments, the C₁-C₃hydrocarbon yield (“gaseous hydrocarbon yield”), or portion of the totalcarbon in the CO in the FT feed provided to an FT reactor, which isconverted to C₁-C₃ hydrocarbons in the FT product removed from thereactor, is less than about 30% (e.g., from about 1% to about 30%) oreven less than about 20% (e.g., from about 3% to about 20%). Asdescribed above with respect to conversion, amounts provided to andremoved from the reactor may be expressed in terms of flow rates.

Embodiments of the invention are therefore directed to a process forproducing C₄ ⁺ hydrocarbons from a synthesis gas comprising H₂ and CO,for example a synthesis gas product, or an FT feed, as described above.The synthesis gas product or FT feed may generally be produced byreforming (conventional reforming, dry reforming, or CO₂-steamreforming). The process comprises contacting the synthesis gas with anFT catalyst comprising at least about 10 wt-% Co and/or optionally othertransition metal(s) described above, on a solid support, for example arefractory metal oxide such as alumina. The process comprises convertingH₂ and CO in the synthesis gas to hydrocarbons, including C₄ ⁺hydrocarbons, provided in an FT product, for example as describedherein.

Advantageously, in the absence of FT product recycle, compression costsare saved and the overall design of the integrated process issimplified. To the extent that this requires an increase in the per-passconversion and associated shift in the distribution of hydrocarbons inthe FT product toward those having increased numbers of carbon atoms,including normal C₂₀ ⁺ hydrocarbons that are undesirable, it should beappreciated that aspects of the invention are associated with thediscovery of important, further downstream processing strategies forconverting these normal C₂₀ ⁺ hydrocarbons to normal and/or branchedC₄-C₁₉ hydrocarbons, which contribute to the yield of desired naphthaboiling-range hydrocarbons, jet fuel boiling-range hydrocarbons, and/ordiesel boiling-range hydrocarbons. An optional further downstreamprocessing stage, namely a finishing stage for carrying out thisconversion, is described below.

Finishing Stage

An optional finishing stage may be desirable, as described above, inembodiments in which the C₄ ⁺ hydrocarbons in the FT product includenormal C₂₀ ⁺ hydrocarbons. In particular, a wax fraction of the C₄ ⁺hydrocarbons may comprise such high carbon number hydrocarbons, withthis wax fraction referring to hydrocarbons that are solid at roomtemperature and that not only represent a loss in yield of hydrocarbonshaving greater utility as liquid fuels, but also pose significantproblems in terms of causing detrimental wax accumulation within processpiping, in addition to difficulties associated with transporting andblending.

In the finishing reactor(s) of a finishing stage, at least a portion ofthe normal C₂₀ ⁺ hydrocarbons in the FT product are converted to normaland/or branched C₄-C₁₉ hydrocarbons, according to hydroisomerization andhydrocracking reactions occurring in the reactor(s). In particular, afinishing feed may comprise some or all of the FT product, optionallyfollowing one or more intervening operations such as cooling, heating,pressurizing, depressurizing, separation of one or more components,addition of one or more components, and/or reaction of one or morecomponents. In view of the temperatures and pressures typically used inthe finishing reactor(s) of the finishing stage relative to those usedin the FT reactor(s) of the FT synthesis stage, the FT product may beheated, prior to conversion of normal C₂₀ ⁺ hydrocarbons in the FTproduct in the finishing stage, to a temperature suitable for afinishing reactor used in this stage, as described herein. In someembodiments, this heating may be the only intervening operation to whichthe FT product is subjected, to provide the finishing feed.Alternatively, for even greater operational simplicity and efficiency,even this heating may be omitted, in view of the possibility for the FTreaction conditions to include a temperature that is the same orsubstantially the same as (e.g., within about 10° C. (18° F.) of) thatused in the downstream finishing stage, for example within a temperaturerange as described below with respect to the finishing reactionconditions. In other embodiments, intervening operations that may beomitted include pressurizing and depressurizing, as it has beendiscovered that finishing reaction conditions can advantageously includea same or substantially same pressure as described above with respect toFT reaction conditions. For example, a pressure in a finishing reactorcan be the same pressure as in an upstream FT reactor, reduced by anominal pressure drop associated with the piping and possibly otherprocess equipment between these reactors. Therefore, costs forpressurization (compression) or depressurization (expansion) of the FTproduct, upstream of the finishing reactor, can be advantageouslyavoided. As with intervening operations between the reforming stage andFT synthesis stage, the use of no intervening operations, limitedintervening operations, and/or the omission of certain interveningoperations between the FT synthesis stage and finishing stage results inadvantages associated with the overall simplification of the integratedprocess. Particular advantages result, for example, if all orsubstantially all of the synthesis gas product is used in the FT feedand/or all or substantially all of the FT product is used in thefinishing feed. In other embodiments, all or substantially all of thesynthesis gas product, except for a condensed water-containing portion,is used in the FT feed and/or all or substantially all of the FT productis used in the finishing feed.

Conditions in the finishing reactor(s) are suitable for the conversionof normal C₂₀ ⁺ hydrocarbons to C₄-C₁₉ hydrocarbons, according tofinishing reactions that include or possibly consist ofhydroisomerization and/or hydrocracking reactions. A finishing reactormay be incorporated into an FT reactor, for example by using a bed offinishing catalyst directly following a bed of FT catalyst within asingle vessel, or otherwise interspersing the two catalyst types withina single vessel. However, generally the use of at least one separatefinishing reactor (e.g., in a separate finishing reactor vessel) ispreferred, such that finishing reaction conditions can be maintainedindependently of FT reaction conditions as described above. A separatefinishing reactor may be advantageous, for example, for (i) maintainingthe finishing catalyst in a different reactor type, compared to the FTreactor, such as maintaining the finishing catalyst in a fixed bedreactor that is normally simpler in design compared to the FT reactor,as a fixed bed reactor normally does involve not the same designconstraints in terms of the ability to remove reaction heat, (ii)removing and/or replacing the finishing catalyst at times that do notnecessarily coincide with (e.g., at differing intervals relative to)removing and/or replacing the FT catalyst, and/or (iii) operating thefinishing reactor at a different temperature (e.g., at a highertemperature) compared to the FT reactor. With respect to the use of aseparate finishing reactor, it may be important to maintain the FTproduct (or at least any portion of this product used in the finishingreactor), from the outlet (effluent) of the FT reactor to the inlet ofthe finishing reactor, at an elevated temperature to avoid deposition ofany normal C₂₀ ⁺ hydrocarbons, and other hydrocarbons having similarlyhigh melting temperatures, as solid wax. Such deposition can result notonly in losses of desired product that would otherwise be produced fromconversion in the finishing stage, but also in the plugging and/orfouling of process equipment, leading to operational failure. The use ofa finishing reactor may also be simplified if condensation of any normalC₂₀ ⁺ hydrocarbons is avoided, i.e., if all or substantially all of theFT product is maintained in the vapor phase from the outlet of the FTreactor to the inlet of the finishing reactor. For example, to avoiddeposition and/or condensation, the FT product may be maintained at atemperature of at least about 66° C. (150° F.), at least about 121° C.(250° F.), at least about 216° C. (420° F.), or even at least about 327°C. (620° F.), from the effluent of the FT reactor to the inlet of thefinishing reactor, such as in the case of heating the FT product fromthis temperature to a temperature suitable for a finishing reactor, asdescribed herein.

In representative embodiments, finishing reaction conditions (suitablefor use in at least one finishing reactor) can include a temperature ina range from about 232° C. (450° F.) to about 399° C. (750° F.), or fromabout 304° C. (580° F.) to about 371° C. (700° F.). Other finishingreaction conditions can include a gauge pressure from about 621 kPa (90psig) to about 3.38 MPa (490 psig), or from about 2.00 MPa (290 psig) toabout 3.10 MPa (450 psig).

In the finishing reactor(s), the finishing feed may be contacted with asuitable finishing catalyst (e.g., bed of finishing catalyst particlesdisposed within the finishing reactor) under finishing reactionconditions, which may include the temperatures and/or pressuresdescribed above. As also described above, the finishing catalystpreferably has activity for hydrocracking and/or hydroisomerization ofnormal C₂₀ ⁺ hydrocarbons present in the FT product. These hydrocarbons,characteristic of solid wax, result from the carbon number distributionof normal hydrocarbons produced by the Fischer-Tropsch reactionchemistry, in conjunction with C₄-C₁₉ hydrocarbons that are moredesirable as components of liquid fuels, as described herein. As isunderstood in the art, hydroisomerization refers to reactions of normalhydrocarbons in the presence of hydrogen to produce branchedhydrocarbons. Hydrocracking refers to reactions of hydrocarbons withhydrogen to produce hydrocarbons having a lower number of carbon atomsand consequently a lower molecular weight. Hydroisomerization isbeneficial for improving characteristics of hydrocarbons having a lowernumber of carbon atoms (e.g., C₄-C₁₉ hydrocarbons) and useful ascomponents of liquid fuels, which hydrocarbons may be present in thefinishing feed and/or FT product or which may be produced byhydrocracking in the finishing reactor(s). These characteristics includea higher octane number (e.g., research octane number and/or motor octanenumber) of naphtha boiling-range hydrocarbons present in the finishingproduct, relative to that of the finishing feed and/or FT product. Thesecharacteristics also include a reduced pour point of dieselboiling-range hydrocarbons present in the finishing product, relative tothat of the finishing feed and/or FT product. Hydrocracking isbeneficial for its overall impact on the carbon number distribution ofthe finishing feed, which may correspond to that of the FT product, andin particular for reducing the percentage by weight of, and possiblyeliminating, normal C₂₀ ⁺ hydrocarbons present in finishing feed and/orFT product. These hydrocarbons, being solid at room temperature, hinderthe ability of products containing such hydrocarbons to be transportedvia a normal pipeline.

As both hydroisomerization and hydrocracking reactions require hydrogen,in preferred embodiments this hydrogen is present in the finishing feedand/or FT product to the finishing reactor. For example, hydrogen in thesynthesis gas product that is unconverted in the downstream FT reactormay allow operation of the finishing reactor without the need for asupplemental source of hydrogen being added to the finishing reactor ordownstream of the FT reactor. According to some embodiments, hydrogen ispresent in the finishing feed and/or FT product at a concentration ofleast about 20 mol-% (e.g., from about 20 mol-% to about 75 mol-%), atleast about 30 mol-% (e.g., from about 30 mol-% to about 65 mol-%), orat least about 40 mol-% (e.g., from about 40 mol-% to about 60 mol-%),without the introduction of a supplemental source of hydrogen, beyondthe hydrogen produced in the reforming stage and/or present in thesynthesis gas product. According to other embodiments, a supplementalsource of hydrogen, added to a finishing reactor, or upstream of afinishing reactor, of the finishing stage (e.g., downstream of an FTreactor of the FT synthesis stage), may be used to achieve such hydrogenconcentrations. A representative supplemental source of hydrogen ishydrogen that has been purified (e.g., by PSA or membrane separation) orhydrogen that is impure (e.g., syngas).

Representative finishing catalysts, to the extent that they haveactivity for converting wax, i.e., hydroisomerization and hydrocrackingactivity with respect to normal C₂₀ ⁺ hydrocarbons as described above,may also be referred to as dewaxing catalysts. Examples of finishing ordewaxing catalysts comprise at least one dewaxing active (e.g.,hydroisomerization and/or hydrocracking active) metal on a solidsupport. The phrase “on a solid support” is intended to encompasscatalysts in which the active metal(s) is/are on the support surfaceand/or within a porous internal structure of the support. Representativedewaxing active metals may be selected from the Groups 12-14 of thePeriodic Table, such as from Group 13 or Group 14 of the Periodic Table.A particular dewaxing active metal is gallium. The at least one dewaxingactive metal may be present in an amount, for example, from about 0.1wt-% to about 3 wt-%, or from about 0.5 wt-% to about 2 wt-%, based onthe weight of the dewaxing catalyst. If a combination of dewaxing activemetals are used, such as a combination of metals selected from Groups12-14 of the Periodic Table, then such metals may be present in acombined amount within these ranges. Generally, the dewaxing catalystsmay comprise no metal(s) on the support in an amount, or combinedamount, of greater than about 1 wt-%, or greater than about 0.5 wt-%,based on the weight of the dewaxing catalyst, other than the dewaxingactive metal(s) described above (e.g., no metals other than metals ofGroups 12-14 of the Periodic Table, no metals other than metals ofGroups 13 or Group 14 of the Periodic Table, or no metals other thangallium, in this amount or combined amount). Preferably, the dewaxingcatalyst comprises no metals on the support, other than the dewaxingactive metal(s) described above (e.g., no metals other than metals ofGroups 12-14 of the Periodic Table, no metals other than metals ofGroups 13 or Group 14 of the Periodic Table, or no metals other thangallium).

In order to promote hydrocracking activity, the solid support of thefinishing catalyst or dewaxing catalyst may be more particularly a solidacidic support. The acidity of a support may be determined, for example,by temperature programmed desorption (TPD) of a quantity of ammonia(ammonia TPD), from an ammonia-saturated sample of the support, over atemperature from 275° C. (527° F.) to 500° C. (932° F.), which is beyondthe temperature at which the ammonia is physisorbed. The quantity ofacid sites, in units of millimoles of acid sites per gram (mmol/g) ofsupport, therefore corresponds to the number of millimoles of ammoniathat is desorbed per gram of support in this temperature range. Arepresentative solid support comprises a zeolitic or non-zeoliticmolecular sieve and has at least about 15 mmol/g (e.g., from about 15 toabout 75 mmol/g) of acid sites, or at least about 25 mmol/g (e.g., fromabout 25 to about 65 mmol/g) of acid sites, measured by ammonia TPD. Inthe case of zeolitic molecular sieves, acidity is a function of thesilica to alumina (SiO₂/Al₂O₃) molar framework ratio, and, inembodiments in which the solid support comprises a zeolitic molecularsieve (zeolite), its silica to alumina molar framework ratio may be lessthan about 60 (e.g., from about 1 to about 60), or less than about 40(e.g., from about 5 to about 40). Particular solid supports may compriseone or more zeolitic molecular sieves (zeolites) having a structure typeselected from the group consisting of FAU, FER, MEL, MTW, MWW, MOR, BEA,LTL, MFI, LTA, EMT, ERI, MAZ, MEI, and TON, and preferably selected fromone or more of FAU, FER, MWW, MOR, BEA, LTL, and MFI. The structures ofzeolites having these and other structure types are described, andfurther references are provided, in Meier, W. M, et al., Atlas ofZeolite Structure Types, 4^(th) Ed., Elsevier: Boston (1996). Specificexamples include zeolite Y (FAU structure), zeolite X (FAU structure),MCM-22 (MWW structure), and ZSM-5 (MFI structure), with ZSM-5 beingexemplary.

Solid supports other than zeolitic and non-zeolitic molecular sievesinclude metal oxides, such as any one or more of silica, alumina,titania, zirconia, magnesium oxide, calcium oxide, strontium oxide, etc.In representative embodiments, the solid support may comprise (i) asingle type of zeolitic molecular sieve, (ii) a single type ofnon-zeolitic molecular sieve, or (iii) a single type of metal oxide,wherein (i), (ii), or (iii) is present in an amount greater than about75 wt-% (e.g., from about 75 wt-% to about 99.9 wt-%) or greater thanabout 90 wt-% (e.g., from about 90 wt-% to about 99 wt-%), based on theweight of the dewaxing catalyst. Other components of the support, suchas binders and other additives, may be present in minor amounts, such asin an amount, or combined amount, of less than about 10 wt-% (e.g., fromabout 1 wt-% to about 10 wt-%), based on the weight of the dewaxingcatalyst.

An exemplary dewaxing catalyst comprises gallium as the dewaxing activemetal, present in an amount as described above (e.g., from about 0.5wt-% to about 2 wt-%, such as about 1 wt-%, based on the weight of thedewaxing catalyst) on a support comprising, or possibly consistingessentially of, ZSM-5. Representative silica to alumina molar frameworkratios of the ZSM-5 are describe above.

Finishing or dewaxing catalysts and finishing reaction conditionsdescribed herein are generally suitable for achieving a conversion ofnormal C₂₀ ⁺ hydrocarbons (e.g., normal C₂₀-C₆₀ hydrocarbons) of atleast about 80% (e.g., from about 80% to about 100%), at least about 85%(e.g., from about 85% to about 98%), or at least about 90% (e.g., fromabout 90% to about 95%). Such high conversion levels are important forimproving the quality of the FT product, especially in terms of itsability to be transportable (e.g., via pipeline) as a liquid fuel,without the need for separation or conversion of solid wax. Theconversion of normal C₂₀ ⁺ hydrocarbons to lower molecular weight,C₄-C₁₉ hydrocarbons also improves the overall yield of thesehydrocarbons, compared to the operation of the FT synthesis stage inisolation. Preferably, in the finishing stage (e.g., in a finishingreactor of this stage), at least about 75% (e.g., from about 75% toabout 100%), at least about 85% (e.g., from about 85% to about 98%), orat least about 90% (e.g., from about 90% to about 97%) of the normal C₂₀⁺ hydrocarbons in the FT product are converted to C₄-C₁₉ hydrocarbons.That is, the yields of C₄-C₁₉ hydrocarbons from the conversion of normalC₂₀ ⁺ hydrocarbons in the finishing stage are within these ranges.Preferably, the finishing product (or hydroisomerization/hydrocrackingproduct of the finishing reactor) comprises less than about 2 wt-%, oreven less than about 1 wt-% of hydrocarbons that are solid at roomtemperature (e.g., normal C₂₀ ⁺ hydrocarbons). In representativeembodiments, normal C₂₀ ⁺ hydrocarbons are converted (e.g., at completeor substantially complete conversion and/or within the conversion rangesgiven above) in the finishing stage (e.g., in at least one finishingreactor of this stage), with a yield of (i) isoparaffinic (branched)hydrocarbons from about 25% to about 70%, or from about 40% to about60%, (ii) aromatic hydrocarbons from about 10% to about 35% or fromabout 15% to about 25%, (iii) gasoline boiling-range hydrocarbons fromabout 50% to about 95% or from about 70% to about 90%, (iv) dieselboiling-range hydrocarbons from about 5% to about 45% or from about 10%to about 30%, and/or (v) VGO boiling-range hydrocarbons of less thanabout 1% or less than about 0.5%, with these yields referring to thepercentage of the total carbon in the normal C₂₀ ⁺ hydrocarbons in thefinishing feed provided to a finishing reactor, which is converted tothese components in the finishing product. Advantageously, isoparaffinichydrocarbons improve the quality of diesel boiling-range hydrocarbons byreducing both the pour point and the cloud point of this fraction. Bothisoparaffinic hydrocarbons and aromatic hydrocarbons improve the qualityof gasoline boiling-range hydrocarbons by increasing the octane number(e.g., research octane number and/or motor octane number) of thisfraction. In representative embodiments, the gasoline boiling-rangehydrocarbons obtained from conversion of normal C₂₀ ⁺ hydrocarbons inthe finishing stage have a research octane number of at least about 75(e.g., from about 75 to about 85).

As described above, conversion levels of normal C₂₀ ⁺ hydrocarbons inthe finishing stage (e.g., in the at least one finishing reactor of thisstage) may be below 100% and therefore allow for a portion of thesenormal C₂₀ ⁺ hydrocarbons in the finishing feed to remain unconverted.To achieve complete conversion of normal C₂₀ ⁺ hydrocarbons, such ascomplete conversion to C₄-C₁₉ hydrocarbons and/or branched C₂₀ ⁺hydrocarbons, finishing reaction conditions may be made more severe,such as by increasing temperature, increasing pressure, and/ordecreasing WHSV.

However, it is to be understood that complete conversion of normal C₂₀ ⁺hydrocarbons is not a requirement to achieve complete “dewaxing” of theFT product and/or finishing feed, in the sense of providing a finishingproduct that is free of solid phase hydrocarbons and therefore easilytransportable as a liquid fuel, according to preferred embodiments.Incomplete conversion of normal C₂₀ ⁺ hydrocarbons (such as achievingconversion levels within certain ranges described above) can nonethelessprovide a finishing product in which sufficient products resulting fromthe conversion of normal C₂₀ ⁺ hydrocarbons, namely (i) sufficientnon-normal C₂₀ ⁺ hydrocarbons (e.g., branched C₂₀ ⁺ hydrocarbons) havingmelting points below room temperature (20° C.) and/or (ii) sufficientC₄-C₁₉ hydrocarbons, are present in the finishing product, to the extentthat any unconverted normal C₂₀ ⁺ hydrocarbons are dissolved at roomtemperature in the finishing product comprising (i) and (ii).

Embodiments of the invention are therefore directed to the use of afinishing stage, following an FT synthesis stage, to improve the overallselectivities to, and yields of, desired products and/or decrease theoverall selectivities to, and yields of, undesired products(particularly wax), relative to the FT synthesis stage in the absence ofthe finishing stage (i.e., relative to a baseline FT synthesis stage orFT synthesis reaction). For example, the finishing stage canbeneficially convert some or all wax (e.g., at the conversion levels ofnormal C₂₀ ⁺ hydrocarbons as described above) produced by the FTsynthesis reaction, thereby decreasing the selectivity to (and/or yieldof) wax, in the combined FT synthesis and finishing stages relative tothe baseline FT synthesis stage. In representative embodiments, theselectivity to (and/or yield of) wax is decreased from a value fromabout 10% to about 50%, such as from about 20% to about 45%, in thebaseline FT synthesis stage to a value from about 0% to about 10%, suchas from about 0.5% to about 5%, in the combined FT synthesis andfinishing stages. Preferably, this selectivity to (and/or yield of) waxis decreased to less than about 0.5%. As described above, smallquantities of wax in the finishing product can be acceptable to theextent that any unconverted normal C₂₀ ⁺ hydrocarbons, and/or anyhydrocarbons generally that melt above room temperature, are present inan amount that is below their solubility in the finishing product (i.e.,in an amount such that they may be completely dissolved in the finishingproduct). In other representative embodiments, the selectivity to(and/or yield of) of C₄-C₁₉ liquid hydrocarbons is increased from avalue from about 15% to about 45%, such as from about 20% to about 35%,in the baseline FT synthesis stage to a value from about 40% to about75%, such as from about 50% to about 70%, in the combined FT synthesisand finishing stages. Selectivities to wax or C₄-C₁₉ hydrocarbons, withrespect to the baseline FT synthesis stage and combined FT synthesis andfinishing stages, are based on the percentage of carbon in CO convertedby FT, which results in wax or C₄-C₁₉ liquid hydrocarbons, respectively.Yields of wax or C₄-C₁₉ hydrocarbons, with respect to the baseline FTsynthesis stage and combined FT synthesis and finishing stages, arebased on the percentage of carbon in CO introduced to the FT synthesisstage (e.g., CO introduced with the FT feed, whether converted orunconverted), which results in wax or C₄-C₁₉ liquid hydrocarbons,respectively. These (i) decreases in selectivity to (and/or yield of)wax, and/or (ii) increases in selectivity to (and/or yield of) C₄-C₁₉liquid hydrocarbons, as a result of incorporating the finishing stage(e.g., finishing reactor), can be achieved without a significantdifference between the CO conversion obtained in the baseline FTsynthesis stage and that obtained in the combined FT synthesis andfinishing stages. For example, the CO conversion values obtained in boththe baseline FT synthesis stage and combined FT synthesis and finishingstages may be within a range as described above with respect to theperformance criteria of the FT synthesis stage. That is, the finishingstage typically does not significantly impact the CO conversion obtainedin the FT synthesis stage alone, such that the CO conversion achieved inboth the baseline FT synthesis stage and combined FT synthesis andfinishing stages may be the same or substantially the same.

The conversion levels in the finishing stage, as described above, may bebased on “per-pass” conversion, achieved in a single pass through thefinishing stage (e.g., a finishing reactor of this stage), or otherwisebased on overall conversion, achieved by returning a recycle portion ofthe finishing product back to the finishing stage (e.g., a finishingreactor of this stage), as described above with respect to the FTsynthesis stage. A desired conversion of normal C₂₀ ⁺ hydrocarbons maybe achieved by adjusting the finishing reaction conditions describedabove (e.g., finishing reaction temperature and/or pressure), and/oradjusting the weight hourly space velocity (WHSV), as defined above. Thefinishing reaction conditions may include a weight hourly space velocity(WHSV) generally from about 0.05 hr⁻¹ to about 35 hr⁻¹, typically fromabout 0.1 hr⁻¹ to about 20 hr⁻¹, and often from about 0.5 hr⁻¹ to about10 hr⁻¹. The finishing reaction conditions may optionally includereturning a recycle portion of the finishing product, exiting thefinishing reactor, back to the finishing feed for combining with thefinishing feed, or otherwise back to the finishing reactor itself.Recycle operation allows for operation at relatively low “per-pass”conversion through the finishing reactor, while achieving a high overallconversion due to the recycle. Preferably, however, the finishingreaction conditions include little or even no finishing product recycle.For example, the finishing reaction conditions may include a weightratio of recycled finishing product to finishing feed (i.e., a “recycleratio”), with this recycled finishing product and finishing feedtogether providing a combined feed to the FT reactor, of those describedabove with respect to the FT synthesis stage. Preferably, the recycleratio may be 0, meaning that no finishing product recycle is used, suchthat the per-pass conversion is equal to the overall conversion.Advantageously, in the absence of finishing product recycle, utilitycosts are saved and the overall design of the integrated process issimplified.

Embodiments of the invention are therefore directed to a process forconverting C₂₀ ⁺ hydrocarbons (e.g., normal C₂₀ ⁺ hydrocarbons) in afeed comprising C₄ ⁺ hydrocarbons, such as a finishing feed as describedabove, which may comprise all or a portion of an FT product as describedabove. The feed comprising C₄ ⁺ hydrocarbons may comprise, for example,C₂₀ ⁺ hydrocarbons in an amount of at least about 5 wt-% (e.g., fromabout 5 wt-% to about 30 wt-%), or at least about 10 wt-% (e.g., fromabout 10 wt-% to about 25 wt-%), based on the weight of totalhydrocarbons, or based on the weight of the feed. The feed may furthercomprise hydrogen (e.g., in an amount as described above with respect toa finishing feed), CO, and/or CO₂. The process comprises contacting thefeed with a finishing or dewaxing catalyst as described above, forexample comprising an active metal selected from Groups 12-14 of thePeriodic Table (e.g., gallium) on a zeolitic molecular sieve support(e.g., ZSM-5), to achieve conversion of the C₂₀ ⁺ hydrocarbons atconversion levels, and with yields and selectivities to lower numberhydrocarbons, and hydrocarbon fractions, as well as other performancecriteria, as described herein.

Overall Performance Criteria, Advantages, and Exemplary Embodiments

An integrated process as described above, and particularly utilizing thecombination of (i) a dry reforming or CO₂-steam reforming process asdescribed above, in combination with (ii) Fischer-Tropsch synthesis, and(iii) optional finishing (dewaxing), may be referred to as an“integrated CSR-FT process,” and used for the direct conversion ofhydrocarbons such as methane in natural gas to one or more liquid fuels.Such liquid fuel(s) may be provided in a finishing product exiting thefinishing stage (e.g., a reactor of this stage) as described above,together with low carbon number hydrocarbons, such as C₁-C₃hydrocarbons. These low carbon number hydrocarbons, together withresidual, unconverted gases (e.g., H₂, CO, and/or CO₂) may be separatedfrom the liquid fuel(s) (e.g., comprising C₄-C₁₉ hydrocarbons andoptionally branched C₂₀ ⁺ hydrocarbons) using a flash separation vesselproviding a vapor-liquid equilibrium separation stage. Alternatively,multiple vapor-liquid equilibrium separation stages may be used, as inthe case of separation using distillation, to separate such low carbonnumber hydrocarbons and also separate the liquid fuels, for example byseparating a fraction comprising predominantly, substantially all, orall, gasoline boiling-range hydrocarbons from a fraction comprisingpredominantly, substantially all, or all, diesel boiling-rangehydrocarbons. In yet other embodiments, a flash separation vessel may beused to perform an initial separation of low carbon number hydrocarbonsand residual gases from the finishing product, followed by separation ofliquid fuels in the finishing product using distillation.

A number of advantages arise in integrated CSR-FT processes describedherein, which include those associated with operation of the FTsynthesis stage at a high per-pass conversion, as described above. Theseadvantages include an option to operate the FT synthesis stage withoutrecycle and with a shift in the distribution of hydrocarbons in the FTproduct toward those having higher numbers of carbon atoms and presentin liquid fuels, thereby decreasing the yield of less desirable C₁-C₃hydrocarbons. In representative embodiments, integrated CSR-FT processescan convert hydrocarbons (e.g., methane) present in a gaseous mixtureand/or hydrocarbon-containing feedstock as described above and fed tothe process, such that at least about 70% (e.g., from about 70% to about95%), or at least about 85% (e.g., from about 85% to about 95%) of thecarbon, initially present in hydrocarbons converted in the process, ispresent in C₄-C₁₉ liquid hydrocarbons in the finishing product. That is,the selectivity of the overall integrated CSR-FT process to liquidfuel(s) comprising these hydrocarbons (e.g., naphtha boiling-rangehydrocarbons and diesel boiling-range hydrocarbons) may be in theseranges. Also, at most about 25% (e.g., from about 5% to about 25%), orat most about 15% (e.g., from about 10% to about 15%) of the carbon,initially present in hydrocarbons converted in the process, may bepresent in C₁-C₃ hydrocarbons in the finishing product. That is, theselectivity of the overall integrated CSR-FT process to these low carbonnumber hydrocarbons may be in these ranges. In addition, to the extentthat these low carbon number hydrocarbons may be separated as a vaporfraction of the finishing product, this vapor fraction, due to itscombustive heating (fuel) value, may be combusted to provide heat energyelsewhere in the integrated CSR-FT process, particularly in the furnaceor hotbox of a reforming reactor of the reforming stage. This wouldallow for the generation of at least a portion, and possibly all, of theheat needed to sustain the endothermic dry reforming and/or CO₂-steamreforming reactions of the reforming stage, particularly in view of thefact that the vapor fraction typically comprises not only C₁-C₃hydrocarbons, but also residual H₂ and/or CO that are likewisecombustible.

Moreover, the use of the optional finishing stage can effectivelyconvert all or substantially all wax (e.g., comprising normal C₂₀ ⁺hydrocarbons) to hydrocarbons having lower carbon numbers (e.g., withinthe range of C₄-C₁₉ hydrocarbons) and useful as liquid fuels. Theoptional finishing stage can also convert a portion of the wax toisoparaffinic C₂₀ ⁺ hydrocarbons having a melting point below roomtemperature. To the extent that any hydrocarbons having a melting pointabove room temperature are present in the finishing product, the amountof such hydrocarbons may be sufficiently small so as to be completelysoluble in this product, thereby beneficially rendering a liquidfraction of the finishing product suitable for transport via pipeline.Furthermore, the finishing stage can isomerize other hydrocarbons (e.g.,C₄-C₁₉ hydrocarbons) present in the FT product and/or finishing feed,thereby increasing the octane number of gasoline boiling-rangehydrocarbons and/or decreasing the pour point and/or cloud point ofdiesel boiling-range hydrocarbons present in the finishing product,relative to the respective values in the FT product and/or finishingfeed.

FIG. 3 depicts a flowscheme of a representative, integrated CSR-FTprocess 100, in which a dry reforming or CO₂-steam reforming process 10,such as described above and depicted in FIG. 1A or 1 i, is integratedwith downstream processing steps using FT reactor 20 and finishingreactor 30, for producing liquid hydrocarbons as described above.According to integrated CSR-FT process 100, gaseous mixture 4 may beprovided via a connection, such as from system input 15, to a source ofthe gaseous mixture or a source of one or more components of thisgaseous mixture (e.g., a hydrocarbon-containing feedstock such asnatural gas), as described above. From system input 15, gaseous mixture4 may be directed to reforming reactor 5, which may operate underreforming conditions as described above and may optionally comprisereforming catalyst 6, such as a catalyst as described above. Synthesisgas product 7, received from reforming reactor 5, may be directed tosynthesis gas product cooler 17 and cooled, for example, from atemperature representative of a reforming condition as described above,to a temperature representative of a downstream FT reaction condition asdescribed above. Cooled synthesis gas product 19 may be received fromsynthesis gas product cooler 17 and directed to optional condenser 21,for the removal of condensed water 22 from cooled synthesis gas product19. Condensed water 22 in this case may be provided as a system water(or aqueous product) output.

Whether or not optional condenser 21 is included or excluded fromintegrated CSR-FT process 100, cooled synthesis gas product 19 may bedirected to compressor 23 to increase the pressure of cooled synthesisgas product 19 to a pressure representative of an FT reaction conditionas described above. FT feed 27 may be received from compressor 23 anddirected to FT reactor 20, which may operate under FT reactionconditions as described above and may optionally comprise an FT catalystas described above. Therefore, all or part of synthesis gas product 7may be directed to FT reactor 20, to form all or part of FT feed 27(e.g., a part of synthesis gas product 7, obtained after condensingwater, may form all, or substantially all, of FT feed 27). FT product 29may be received from FT reactor 20 and directed to optional FT productheater 31. Optional FT product heater 31 may be used to heat FT product29 to a temperature representative of a finishing reaction condition asdescribed above. Alternatively, both FT reactor 20 and downstreamfinishing reactor 30 may be operated at the same or substantially thesame temperature, such that optional FT product heater 31 may beexcluded from integrated CSR-FT process 100. All or part of FT product29 may be directed to finishing reactor 30, to form all or part offinishing feed 32 (e.g., all of FT product 29 may form all, orsubstantially all, of finishing feed 32). Finishing reactor 30 mayoperate under finishing reaction conditions as described above and mayoptionally comprise a finishing catalyst as described above. Finishingproduct 33 may be received from finishing reactor 30 and directed tofinishing product separator 50 that provides separated fractions offinishing product 33, such as vapor fraction 37 and liquid fraction 39,to a system vapor output 40 and to a system liquid output 45,respectively.

According to alternative embodiments, vapor fraction 37, received fromfinishing product separator 50, may be maintained within integratedCSR-FT process 100 and directed to a furnace or hotbox of reformingreactor 5, as a source of fuel to maintain reforming catalyst 6 at atemperature representative of a reforming condition as described above.In such embodiments, a flue gas effluent (not shown) may be provided asa system vapor output, in lieu of vapor fraction 37. According to otheralternative embodiments, in addition to vapor fraction 37 (which mayalternatively be used as a fuel source for heating reforming reactor 5as described above), separator 50 may provide more defined liquidfractions of finishing product, such as gasoline boiling-rangehydrocarbon containing fraction 41 and diesel boiling-range hydrocarboncontaining fraction 43 as system liquid outputs, for example in the caseof separator 50 operating as a distillation column to resolve thesefractions, as opposed to a single stage (vapor/liquid) flash separator.In this case, liquid fraction 39 may be, more particularly, a highcarbon number hydrocarbon containing fraction, such as a VGOboiling-range containing hydrocarbon fraction. According to furtherembodiments, separator 50 may provide all or substantially all of liquidfraction 39 of finishing product 33 to secondary separator 55 to providemore defined liquid fractions 41, 43 as described above with respect toseparator 50. In this case, as depicted in FIG. 3 , secondary separator55 may be outside of integrated CSR-FT process 100 (e.g., may be used ata remote site to resolve liquid fractions), or otherwise may be includedwithin this process.

Aspects of the invention, in addition to integrated CSR-FT processes,therefore also relate to systems or apparatuses for performing suchprocesses, including integrated CSR-FT process 100 as depicted in FIG. 3. Accordingly, particular embodiments of the invention are directed tosystems or apparatuses for producing C₄ ⁺ hydrocarbons, useful as liquidfuels, from methane and/or other light hydrocarbons. The systems orapparatuses may comprise one or more of the following: (i) a reformingreactor 5 configured to connect, via a system input 15, to a source of agaseous mixture 4, for example a source of natural gas comprisingmethane and CO₂. The reforming reactor 5 may contain a reformingcatalyst 4 as described above and/or may be further configured toproduce or provide, from the gaseous mixture 4, a synthesis gas product7 comprising H₂ and CO, for example under reforming conditions asdescribed above; (ii) a synthesis gas product cooler 17 configured toreceive (and/or cool) the synthesis gas product 7 from the reformingreactor 5. The synthesis gas product cooler 17 may be connected to thereforming reactor 5, or may otherwise have an inlet configured forconnection to an outlet of the reforming reactor 5; (iii) a compressor23 configured to receive (and/or compress) a cooled synthesis gasproduct 19 from the synthesis gas product cooler 17. The compressor 23may be connected to the synthesis gas product cooler 17, or mayotherwise have an inlet configured for connection to an outlet of thesynthesis gas product cooler 17; (iv) an FT reactor 20 configured toreceive an FT feed 27 (e.g., as a compressed output) from the compressor23. The FT reactor 20 may contain an FT catalyst as described aboveand/or may be further configured produce or provide, from the FT feed27, an FT product 29 comprising hydrocarbons, including C₄ ⁺hydrocarbons, by conversion of the H₂ and CO in the synthesis gasproduct 7, for example under FT reaction conditions as described above.The FT reactor 20 may be connected to the compressor 23, or mayotherwise have an inlet configured for connection to an outlet of thecompressor 23; (v) a finishing reactor 30 configured to receive afinishing feed 32, either as a heated output from an optional FT productheater 31, or otherwise directly as FT product 29. The finishing reactor30 may contain a finishing catalyst as described above and/or may befurther configured to produce or provide a finishing product 33comprising normal and branched C₄-C₁₉ hydrocarbons, by conversion ofnormal C₂₀ ⁺ hydrocarbons in the FT product 29, for example underfinishing reaction conditions as described above. The finishing reactor30 may be connected to either the FT reactor 20 or the optional FTproduct heater 31, or the finishing reactor 30 may have an inletconfigured for connection to an outlet of either the FT reactor 20 orthe optional FT product heater 31; and (vi) a finishing productseparator 50 configured to receive the finishing product 33 from thefinishing reactor 30 and further configured to provide or separate, viaa system vapor output 40 and a system liquid output 45, vapor and liquidfractions 37, 39, respectively, of the finishing product 33. Thefinishing product separator 50 may be connected to the finishing reactor33 or may have an inlet configured for connection to an outlet of thefinishing reactor 33. The separator 50 may otherwise be configured toprovide more defined liquid fractions 41, 43 of the finishing product33, as described above, as system liquid outputs. The separator 50 mayalternatively be connected, or configured for connection, to secondaryseparator 55 to provide more defined liquid fractions 41, 43, asdescribed above.

Integrated CSR-FT process 100, or associated system or apparatus, mayoptionally further comprise a condenser 21 configured to condense liquidwater from the cooled synthesis gas product 19. In this case, thecompressor 23 is configured to receive the cooled synthesis gas product19 from the condenser 21, following the removal of condensed water 22,which may be provided as a system water (or aqueous product) output. Thecompressor 23 may be connected to the condenser 21, or may otherwisehave an inlet configured for connection to an outlet of the condenser21.

In view of the above description, it can be appreciated that integratedCSR-FT processes, as well as associated systems and apparatuses, canprovide a highly economical manner of converting hydrocarbon-containinggases such as methane to liquid fuels. Each process step, or each systemelement, can be seamlessly integrated with the next step or element.Such integration is possible, advantageously, without the need forcertain conventional steps and associated elements (equipment) and costs(both capital and operating), such as by the omission of one or more ofthe following steps: (i) removal of CO₂ (e.g., using amine scrubbing)from a source of natural gas with a high CO₂ content, (ii) adjustment ofthe molar H₂:CO ratio of the synthesis gas product, upstream of the FTreactor, (iii) separation of solid or condensed liquid wax (e.g.,comprising normal C₂₀ ⁺ hydrocarbons) from the FT product, upstream ofthe finishing reactor (e.g., for processing of the solid wax in aseparate hydrotreating reactor). In fact, CSR-FT processes, as well asassociated systems and apparatuses, as described herein, canadvantageously operate such that no materials are added and/or removedalong the stages of reforming, FT synthesis, and finishing, except forthe addition of gaseous mixture 4 and the removal of fractions offinishing product 33, with the possibility also of removing condensedwater 22 (or aqueous product). In this manner, integrated CSR-FTprocesses, and associated systems and apparatuses, may be streamlinedand simplified, allowing for their operation and implementation withfavorable economics associated with liquid fuel production.

Moreover, this simplicity allows such integrated CSR-FT processes, andassociated systems and apparatuses, to be operable on a small scale andeven transportable in some embodiments, for example by truck, ship,train, or plane. For example, integrated CSR-FT process 100, or theassociated system or apparatus as described above, may be mounted on askid (skid-mounted) for ease of transport to sources of natural gas,sources of other suitable hydrocarbon-containing feedstocks, and/or evensources of CO₂-containing industrial waste gases. For example,integrated CSR-FT process 100 may advantageously be used for convertingflared natural gas to liquid fuels and reducing greenhouse gas (GHG)emissions at well sites. In the case of such a process beingtransportable, a single process, or its associated system or apparatus,could be used for both of these purposes, and/or used with a variety ofother different gaseous mixtures and components of these mixtures (e.g.,hydrocarbon-containing feedstocks), as described above, even if theirsources are at different locations.

Integration with Biomass Hydropyrolysis

As described above, processes for producing renewable hydrocarbon fuelsfrom the hydropyrolysis of biomass can provide gaseous mixturescomprising methane and/or other light hydrocarbons, in combination withCO₂. Therefore, such gaseous mixtures represent potential feeds toCO₂-steam reforming processes, or otherwise integrated CSR-FT processes,as described above, which can be converted to (i) a hydrogen-containingsynthesis gas, in the case of a CO₂-steam reforming process, or (ii)liquid fuels, in the case of an integrated CSR-FT process. With respectto embodiment (i), the hydrogen-containing synthesis gas can be used,optionally following purification to obtain an H₂-enriched portionthereof, as a source of hydrogen that is used to sustain thehydropyrolysis process. With respect to embodiment (ii), the liquidfuels produced from the integrated CSR-FT process can beneficiallyincrease the overall yield of biogenic (renewable) liquid fuels,relative to the yield that may otherwise be obtained from biomasshydropyrolysis. This increase may be relative to a baseline yield in theabsence of using any reaction stage of an integrated CSR-FT process,which corresponds also to the baseline yield obtained using thereforming stage to produce a synthesis gas product, but in the absenceof converting the H₂ and CO in the synthesis gas product to hydrocarbonsusing an FT synthesis stage, as described above. According to someembodiments, the increase in the yield of biogenic liquid fuels may beat least about 25% (e.g., from about 25% to about 60%), or at leastabout 35% (e.g., from about 35% to about 50%).

FIG. 4 depicts a flowscheme in which hydropyrolysis process 200generates gaseous mixture 4, comprising methane and CO₂, as a feed to aCO₂-steam reforming process 10, such as depicted in FIG. 1A or FIG. 1B.According to this embodiment, therefore, CO₂-steam reforming process 10is integrated with a process for producing a renewable hydrocarbon fuelfrom the hydropyrolysis of biomass. Gaseous mixture 4 may comprisemethane and CO₂, as well as other species, in concentrations asdescribed above with respect to “a hydropyrolysis gaseous mixture.” Inaddition to gaseous mixture 4, hydropyrolysis process 200 also generatessubstantially fully deoxygenated hydrocarbon liquid 61 (e.g., having atotal oxygen content of less than about 2 wt-% or less than about 1wt-%), comprising hydrocarbons that may be separated into gasolineboiling-range hydrocarbon containing fraction 41 and dieselboiling-range hydrocarbon containing fraction 43. Hydropyrolysis process200 may further generate aqueous liquid 63, for example obtained byphase separation from substantially fully deoxygenated hydrocarbonliquid 61. As shown, all or a portion of aqueous liquid 63 mayoptionally be combined with gaseous mixture 4, for example to adjust themolar H₂O:CO₂ ratio of gaseous mixture 4 to CO₂-steam reforming process10, to molar ratios as described above. Hydropyrolysis process 200 mayfurther generate solid char 65. These products of hydropyrolysis process200, including gaseous mixture 4, substantially fully deoxygenatedhydrocarbon liquid 61, and aqueous liquid 63 are generated from feeds tohydropyrolysis process 200, including biomass-containing orbiomass-derived feedstock 67 and hydrogen-containing feed gas stream 69.

With respect to biomass-containing or biomass-derived feedstock 67, theterm “biomass” refers to substances derived from organisms living abovethe earth's surface or within the earth's oceans, rivers, and/or lakes.Representative biomass can include any plant material, or mixture ofplant materials, such as a hardwood (e.g., whitewood), a softwood, ahardwood or softwood bark, lignin, algae, and/or lemna (sea weeds).Energy crops, or otherwise agricultural residues (e.g., loggingresidues) or other types of plant wastes or plant-derived wastes, mayalso be used as plant materials. Specific exemplary plant materialsinclude corn fiber, corn stover, and sugar cane bagasse, in addition to“on-purpose” energy crops such as switchgrass, miscanthus, and algae.Short rotation forestry products, such as energy crops, include alder,ash, southern beech, birch, eucalyptus, poplar, willow, paper mulberry,Australian Blackwood, sycamore, and varieties of paulownia elongate.Other examples of suitable biomass include vegetable oils, carbohydrates(e.g., sugars), organic waste materials, such as waste paper,construction, demolition wastes, and biosludge.

A “biomass-containing” feedstock may comprise all or substantially allbiomass, but may also contain non-biological materials (e.g., materialsderived from petroleum, such as plastics, or materials derived fromminerals extracted from the earth, such as metals and metal oxides,including glass). An example of a “biomass-containing” feedstock thatmay comprise one or more non-biological materials is municipal solidwaste (MSW).

“Biomass-derived,” for example when used in the phrase “biomass-derivedfeedstock,” refers to products resulting or obtained from the thermaland/or chemical transformation of biomass, as defined above, orbiomass-containing feedstocks (e.g., MSW). Representativebiomass-derived feedstocks therefore include, but are not limited to,products of pyrolysis (e.g., bio-oils), torrefaction (e.g., torrefiedand optionally densified wood), hydrothermal carbonization (e.g.,biomass that is pretreated and densified by acid hydrolysis in hot,compressed water), and polymerization (e.g., organic polymers derivedfrom plant monomers). Other specific examples of biomass-derivedproducts (e.g., for use as feedstocks) include black liquor, purelignin, and lignin sulfonate. Biomass-derived feedstocks also extend topretreated feedstocks that result or are obtained from thermal and/orchemical transformation, prior to, or upstream of, their use asfeedstocks for a given conversion step (e.g., hydropyrolysis). Specifictypes of pretreating steps that result in biomass-derived productsinclude those involving devolatilization and/or at least somehydropyrolysis of a biomass-containing feedstock. Therefore, certainpretreated feedstocks are also “biomass-derived” feedstocks, whereasother pretreated feedstocks, for example resulting or obtained fromclassification without thermal or chemical transformation, are“biomass-containing” feedstocks, but not “biomass-derived” feedstocks.

It is therefore also possible to feed to hydropyrolysis process 200, inplace of all or a portion of the biomass-containing feedstock, abiomass-derived feedstock, such as a pretreated feedstock that isobtained from a biomass-containing feedstock, after having beendevolatilized and/or partially hydropyrolyzed in a pretreating reactor(pre-reactor), upstream of a hydropyrolysis reactor vessel. Suchpre-reactor thermal and/or chemical transformations of biomass may beaccompanied by other, supplemental transformations, for example toreduce corrosive species content, reduce hydropyrolysis catalyst poisoncontent (e.g., reduced sodium), and/or a reduce hydroconversion catalystpoison content. Devolatilization and/or partial hydropyrolysis ofbiomass or a biomass-containing feedstock in a pre-reactor may becarried out in the presence of a suitable solid bed material, forexample a pretreating catalyst, a sorbent, a heat transfer medium, andmixtures thereof, to aid in effecting such supplemental transformationsand thereby improve the quality of the pretreated feedstock. Suitablesolid bed materials include those having dual or multiple functions. Inthe case of a pretreating catalyst, those having activity forhydroprocessing of the biomass-containing feedstock, described below,are representative.

It is also possible to feed a biomass-containing feedstock that is apretreated feedstock, obtained after having been subjected to apretreating step, for example a physical classification to improve atleast one characteristic, such as a reduced non-biological materialcontent (e.g., content of glass, metals, and metallic oxides, includingall mineral forms), a reduced average particle size, a reduced averageparticle aerodynamic diameter, an increased average particle surfacearea to mass ratio, or a more uniform particle size.

CO₂-steam reforming process 10, as depicted in FIG. 4 , may include areforming reactor 5, containing a reforming catalyst 6, as depicted inFIG. 1A or 1 , with this catalyst having a composition as describedabove. Reforming reactor 5 may operate under reforming conditions asdescribed above, to produce synthesis gas product 7 comprising H₂ andCO. Optional hydrogen purification module 75, for example utilizingpressure-swing adsorption (PSA) or membrane separation, may be used toobtain H₂-enriched portion 71 of synthesis gas product 7, having ahigher concentration of hydrogen relative to this product (e.g., havinga hydrogen concentration of at least about 80 mol-%, such as from about80 mol-% to about 99 mol-%, or at least about 85 mol-%, such as fromabout 85 mol-% to about 98 mol-%). As shown in FIG. 4 , H₂-enrichedportion 71 may be directed back to hydropyrolysis process 200, toprovide at least a portion, and possibly all, of hydrogen-containingfeed gas stream 69. An H₂-depleted portion of synthesis gas product (notshown) may also be obtained from hydrogen purification module 75 andpossibly combusted to provide heat energy for CO₂-steam reformingprocess 10 or for hydropyrolysis process 200. Hydrogen purificationmodule 75 may be used to preferentially separate, into the H₂-depletedportion, any of CO, CO₂, light (C₁-C₃) hydrocarbons, and/or H₂S.

FIG. 5 depicts a flowscheme in which hydropyrolysis process 200generates gaseous mixture 4, comprising methane and CO₂, as in FIG. 4 .According to the embodiment in FIG. 5 , however, gaseous mixture 4 is afeed to integrated CSR-FT process 100, such as depicted in FIG. 3 .therefore, integrated CSR-FT process 100 is in this case furtherintegrated with a process for producing a renewable hydrocarbon fuelfrom the hydropyrolysis of biomass. Products generated fromhydropyrolysis process 200 are as described above with respect to theembodiment of FIG. 4 . These products include (i) gaseous mixture 4,(ii) substantially fully deoxygenated hydrocarbon liquid 61, comprisinghydrocarbons that may be separated into gasoline boiling-rangehydrocarbon containing fraction 41 and diesel boiling-range hydrocarboncontaining fraction 43, (iii) aqueous liquid 63, and (iv) solid char 65.As also described above with respect to the embodiment of FIG. 4 , allor a portion of aqueous liquid 63 may optionally be combined withgaseous mixture 4, for example to adjust the molar H₂O:CO₂ ratio ofgaseous mixture 4. Because integrated CSR-FT process 100, to whichgaseous mixture is directed in the embodiment of FIG. 5 , includes an FTsynthesis stage and optionally the use of an FT catalyst that issusceptible to sulfur poisoning, it may be preferable, according to someembodiments, to treat gaseous mixture 4 to remove H₂S and/or othersulfur-bearing contaminants, prior to (upstream of) integrated CSR-FTprocess 100.

In the embodiment of FIG. 5 , integrated CSR-FT process 100 providesliquid fraction 39 of finishing product 33, as described above withrespect to FIG. 3 . Liquid fraction 39 may advantageously comprise agasoline boiling-range hydrocarbon containing fraction and/or a dieselboiling-range hydrocarbon containing fraction, either or both of whichmay increase the yields of these fractions 41, 43 relative to yieldsobtained from hydropyrolysis process 200 alone (baseline yields obtainedin the absence of integrated CSR-FT process 100), for example accordingto the yield increases described herein. Also according to theembodiment of FIG. 5 , vapor fraction 37 of finishing product 33 (FIG. 3), comprising methane and/or other light hydrocarbons (e.g., C₂-C₃hydrocarbons), in addition to other combustible species such as residualH₂ and/or CO, may optionally be combusted as a source of fuel. Asdepicted in FIG. 5 , a hydrogen production process 300 as describedabove is used to generate purified hydrogen product 79 by steam methanereforming (SMR) of natural gas 77 supplied to this process. Vaporfraction 37 may therefore be used to generate heat for SMR, as depictedin FIG. 5 , and all or a portion of aqueous liquid 63 fromhydropyrolysis process 200 may be used to generate steam for SMR used inhydrogen production process 300. Purified hydrogen product 79 may beused to provide all or a portion of hydrogen-containing feed gas stream69 to hydropyrolysis process 200.

FIG. 6 provides additional details of a hydropyrolysis process 200, forexample as depicted in FIGS. 4 and 5 and used to convertbiomass-containing or biomass-derived feedstock 67 andhydrogen-containing feed gas stream 69 to provide (i) gaseous mixture 4comprising methane and CO₂, (ii) substantially fully deoxygenatedhydrocarbon liquid 61 comprising liquid hydrocarbon-containingfractions, (iii) aqueous liquid 63, and (iv) solid char 65. As depictedin FIG. 6 , hydropyrolysis process 200 may include two stages ofreaction, carried out in first stage hydropyrolysis reactor 81 andsecond stage hydroconversion reactor 83. Hydropyrolysis reactor 81 mayoperate as a catalytic fluidized bed reactor to devolatilize feedstock67 in the presence of stabilizing hydrogen, producing hydropyrolysisreactor effluent 85. Following the removal of solid char 65 fromhydropyrolysis reactor effluent 85 and cooling in first stage effluentcooler 84, hydropyrolysis vapors 87, including a partially deoxygenatedhydropyrolysis product, light hydrocarbons, H₂, CO, CO₂, and H₂O, aredirected to hydroconversion reactor 83. This reactor may operate as afixed bed, for further catalytic hydrodeoxygenation of the partiallydeoxygenated hydropyrolysis product. Hydroconversion reactor effluent 89is then directed to second stage effluent cooler 86, which condensessubstantially fully deoxygenated hydrocarbon liquid 61 and aqueousliquid 63 from hydroconversion reactor effluent 89. In separator 82,these liquid products 61, 63 of hydropyrolysis process 200 may beseparated by organic/aqueous phase separation, with the less densephase, substantially fully deoxygenated hydrocarbon liquid 61, settlingabove the more dense phase, aqueous liquid 63.

Also in separator 82, product vapor fraction 88, comprising lighthydrocarbons, H₂, CO, CO₂, and H₂O, may be separated by vapor/liquidphase separation. Product vapor fraction 88 may be sent to hydrogenpurification module 75, for example utilizing pressure-swing adsorption(PSA) or membrane separation, to separate recycle hydrogen 97, having ahigher concentration of hydrogen relative to product vapor fraction 88,from gaseous mixture 4. Gaseous mixture 4 may therefore have a lowerconcentration of hydrogen relative to product vapor fraction 88, and mayhave other composition characteristics as described above with respectto representative gaseous mixtures generally, and/or with respect to “ahydropyrolysis gaseous mixture” in particular.

Hydrogen purification module 75 may be used to preferentially separate,into gaseous mixture 4, any or all of light (C₁-C₃) hydrocarbons, CO,CO₂, H₂O, and/or H₂S. Recycle hydrogen 97 may have a hydrogenconcentration, for example, of at least about 80 mol %, such as fromabout 80 mol-% to about 99 mol-%, or at least about 85 mol-%, such asfrom about 85 mol-% to about 98 mol-%. Recycle hydrogen 97 may be usedto provide at least a portion, and possibly all, of hydrogen-containingfeed gas stream 69. Optionally, external make-up hydrogen or freshhydrogen 64 may be combined with recycle hydrogen 97 to providehydrogen-containing feed gas stream 69.

FIG. 7 provides additional details of a hydrogen production process 300,for example as depicted in FIG. 5 . As described above, a hydrogenproduction process may convert natural gas 77 to purified hydrogenproduct 79 using stages of steam methane reforming (SMR) 92, water-gasshift (WGS) reaction 94, and pressure-swing adsorption (PSA) 96. In thiscase, SMR can be used to generate SMR synthesis gas 98, and its hydrogencontent can be increased with WGS reaction 94 to provide WGS product 99.PSA 96 is then used to recover purified hydrogen product 79 and rejectnon-hydrogen impurities (e.g., substantially all non-hydrogenimpurities) in hydrogen-depleted PSA tail gas 91. Hydrogen-depleted PSAtail gas 91 generally comprises (i) unconverted methane (due to methane“breakthrough” from SMR 92), (ii) hydrogen that is not recovered inpurified hydrogen product 79 using PSA 96, and (iii) CO₂, as well astypically CO and H₂O. Hydrogen-depleted PSA tail gas 91 may have othercomposition characteristics as described above with respect to gaseousmixtures generally, and/or with respect to “hydrogen-depleted PSA tailgas.”

Normally, hydrogen-depleted PSA tail gas 91 that is obtained as abyproduct from hydrogen production process is combusted to recover itsfuel value. The energy of this combustion can serve as an importantsource of heat for the furnace or hotbox of SMR 92, as this step ofhydrogen production process 300 operates endothermically and at hightemperatures (e.g., as high as 950° C. (1742° F.) or higher). Accordingto the process depicted in FIG. 7 , however, hydrogen-depleted PSA tailgas 91 is directed first to integrated CSR-FT process 100, for exampleas depicted in FIG. 3 and described above. Depending on the compositionof hydrogen-depleted PSA tail gas 91, supplemental hydrocarbon source 95(e.g., natural gas) and/or supplemental steam source 93 may optionallybe combined with hydrogen-depleted PSA tail gas 91 to provide gaseousmixture 4, having a suitable composition as described above. In thismanner, methane and CO₂ from hydrogen-depleted PSA tail gas 91 may beconverted in integrated CSR-FT process 100 to produce liquid fraction 39of finishing product 33 (FIG. 3 ), comprising liquid hydrocarbons usefulas fuels. Although the consumption of methane thereby reduces thecombustive heating value of hydrogen-depleted PSA tail gas 91, the valueof liquid fraction 39 produced outweighs this loss of combustive heatingvalue, which may be replaced, for example using lower cost natural gas.For example, this natural gas, as a supplemental fuel gas (not shown) tothe furnace or hotbox of SMR 92, may be combined with vapor fraction 37of finishing product 33 (FIG. 3 ), as vapor fraction 37 itself canprovide some of the heat needed to maintain SMR 92.

The following examples are set forth as representative of the presentinvention. These examples are not to be construed as limiting the scopeof the invention as other equivalent embodiments will be apparent inview of the present disclosure and appended claims.

Example 1 CO₂-Steam Reforming Studies

Pilot plant scale experiments were performed in which gaseous mixtureswere fed continuously to a CO₂-steam reforming reactor containingcatalyst particles having a composition of 1 wt-% Pt and 1 wt-% Rh on acerium oxide support. The performance of the system for CO₂-steamreforming was tested at conditions of 0.7 hr⁻¹ WHSV, 760° C. (1400° F.),and a gauge pressure ranging from 124 kPa (18 psig) to 172 kPa (25psig). Two types of gaseous mixtures tested were (1) a compositioncontaining methane, ethane, propane, and CO₂, in addition to H₂O, andsimulating that obtained from the combined hydropyrolysis andhydroconversion of biomass (“Renewable Type”), and (2) a typical naturalgas composition having a high level of CO₂ (“Natural Gas Type”). Therenewable type composition provided an example of a methane-containingfeedstock that is also a “hydropyrolysis gaseous mixture,” as describedabove. The natural gas type composition provided an example of amethane-containing feedstock that is also a “natural gas comprisingCO₂,” to which steam, as an H₂O-containing oxidant, has been added, asdescribed above. These gaseous mixtures (combined feeds), and thesynthesis gas products obtained from these feeds, are summarized inTable 1 below.

TABLE 1 CO₂-steam Reforming of Differing Gaseous Mixtures RenewableRenewable Natural Natural gas Type Type Gas Type Type Combined SynthesisCombined Synthesis Feed Gas Product Feed Gas Product methane, mol-% 11.70.3 21.7 .79 ethane, mol-% 5.8 0 5.8 0 propane, mol-% 5.8 0 1.4 0 CO₂,mol-% 23.4 10.6 29.0 8.2 water, mol-% 53.3 12.7 42.1 8.6 H₂, mol-% 51.351.9 CO, mol- % 25.1 30.4 % methane conversion 96 93 % ethane conversion100 100 % propane conversion 100 100 molar H₂:CO ratio 2.05 1.71

From these results, it can be seen that the CO₂-steam reforming catalystand process can provide a synthesis gas product having a molar H₂:COratio that is nearly 2:1 and therefore suitable for subsequent, directprocessing via the Fischer-Tropsch reaction, or at least without a prior(upstream) adjustment of this ratio. Whereas these favorable resultswere obtained at only 760° C. (1400° F.) reaction temperature, lowertemperatures, such as 704° C. (1300° F.) are also possible, in view ofthe high activity of the catalyst. Lower operating temperatures tend toreduce the rate of side reactions that form coke, which deactivates thecatalyst. FIG. 8 illustrates the relationship between temperature andmethane conversion for feeds and reforming catalysts of the type testedin Example 1, and in particular this figure illustrates the ability toachieve greater than 85% methane conversion at 704° C. (1300° F.) andgreater than 95% methane conversion at 760° C. (1400° F.). FIG. 9illustrates how the molar H₂O:CO₂ ratio of the gaseous mixture, forfeeds and reforming catalysts of the type tested in Example 1,influences the molar H₂:CO ratio of the synthesis gas product, attemperatures of both 704° C. (1300° F.) and 760° C. (1400° F.). In viewof the possibility to establish relationships between these parametersfor a given feed, reforming catalyst, and set of operating conditions,the gaseous mixture composition can serve as a convenient control forachieving a target synthesis gas product composition.

Example 2 Sulfur Tolerance of CO₂-Steam Reforming Catalysts

Additional experiments were conducted in which a typical natural gascomposition as described in Example 1 was subjected to CO₂-steamreforming as also described in this example. However, the gaseousmixture or combined feed in this case was spiked with H₂S at aconcentration of 800 mol-ppm. Despite this high level of sulfurcontamination, it was found that the offset in methane conversion waseasily restored by increasing the reforming catalyst bed temperaturefrom 760° C. (1400° F.) to 788° C. (1450° F.). Furthermore, thereforming catalyst surprisingly exhibited long-term stability over 400operating hours (hours on stream) at this temperature, as well as theWHSV and pressure as described above with respect to Example 1. Thisstability, achieved despite the considerable sulfur concentration, wassurprising in view of the sulfur sensitivity of conventional catalystsused for steam methane reforming.

Example 3 Long-Term CO₂-Steam Reforming Testing

The gaseous mixture described in Example 1 as the “Renewable Type” andhaving the composition provided in Table 1 was tested using the catalystand conditions as described in Example 1, to evaluate performance of thesystem for CO₂-steam reforming over an extended period of operation. The“Renewable Type” feed or gaseous mixture also provides an example of arepresentative “hydropyrolysis gaseous mixture” as described above.Long-term stability testing revealed that the composition of thesynthesis gas product obtained was stable over 500 hours of operationunder these constant conditions, demonstrating essentially nodeactivation, over the extended operating period, of the reformingcatalyst. FIG. 10 illustrates the stable synthesis gas productcomposition obtained over this operating period, with a high level ofconversion of methane. FIG. 11 illustrates the stable molar H₂/CO ratioof the synthesis gas product obtained, which was nearly a ratio of 2 andtherefore ideal for use in a downstream FT synthesis reaction to produceliquid hydrocarbons.

Example 4

Evaluation of the Hydroisomerization and Hydrocracking of Wax from FTSynthesis

The FT synthesis reaction typically produces hydrocarbons having a broadrange of molecular weights (and carbon numbers), including normal C₂₀ ⁺hydrocarbons that are solid at room temperature and generally regardedas an undesirable wax product. The use of hydrocracking to eliminatethis wax, by separating it from the FT product and converting it tolower number hydrocarbons, typically adds 1/3 of the capital cost to anFT synthesis complex, as well as a significant amount of complexity.Because it is a solid, wax is not easily shipped through pipelines norblended with crude oil. With the objective of developing a simpleintegrated gas to liquid (GTL) process whereby wax produced in the FTsynthesis reaction could be converted to, and thereby add to the yieldof, (i) lower number hydrocarbons having value as liquid fuels, and/or(ii) isoparaffinic hydrocarbons having melting points below roomtemperature, a simple combined hydroisomerization/hydrocracking reactionfor this purpose was studied. The use of hydroisomerization wasconsidered as a potentially attractive alternative, as this reactionrequires only small amounts of hydrogen. The incorporation of a stepinvolving hydroisomerization directly after the FT synthesis stage, withthis step being provided with all or substantially all of the FT product(e.g., without separation of wax) was therefore proposed as a low costsolution to the problem of wax production in this stage. This step,involving both hydroisomerization and hydrocracking of normal C₂₀ ⁺hydrocarbons, was referred to as the “finishing stage,” utilizing atleast one “finishing reactor.”

In order to investigate possible catalysts for use in thehydroisomerization/hydrocracking of wax, C₂₃-C₆₀ straight chainparaffins were obtained from a commercial supplier of FT wax (Sasol).Batch experiments were performed by adding 200 grams of the wax to astirred Parr bomb reactor. Following this addition of the wax, thetemperature of the reactor was raised under flowing hydrogen or under aflowing synthesis gas (mixture of hydrogen and CO). The reactor, whichhad been loaded with 25 grams of finishing catalyst (orhydroisomerization/hydrocracking catalyst) absolute pressure wasmaintained at 2.76 MPa (400 psia). It was found that a catalystformulation of 1 wt-% gallium on ZSM-5 zeolite support (Ga-ZSM-5catalyst) was effective for converting the wax throughhydroisomerization, combined with hydrocracking. These reactions incombination respectively resulted in the formation of branchedhydrocarbons and also lower molecular weight hydrocarbons, therebyimproving the quality of diesel boiling-range hydrocarbons in terms ofreducing pour point and cloud point, and improving the quality ofgasoline boiling-range hydrocarbons in terms of increasing octanenumber. The results of the batch tests conducted using this catalyst aresummarized in Table 2 below, which includes the recovered productcomposition, following conversion of the wax.

TABLE 2 Conversion of Wax in Batch Testing with Ga-ZSM-5 CatalystTemperature, ° C. 303-342 326-335 299-315 Flowing gas H₂ H₂ + COsynthesis gas H₂ Time of test, min 65 135 210 Wax converted 100% 100%100% Recovered Liquid composition C₃-C₂₆ C₃-C₂₆ C₃-C₂₆ Hydrocarbon Typesparaffins, wt-% 19.3 15.8 16.7 isoparaffins, wt-% 46.3 46.3 53.4naphthenes, wt-% 9.2 8.1 8.9 aromatics, wt-% 17.2 17.7 14.4 olefins,wt-% 7.9 11.8 6.7 Research Octane Number 78.9 79.9 79.7 HydrocarbonBoiling-Range Fractions wt-% gasoline 87.4 84.4 75.1 wt-% jet 10.3 12.121.5 wt-% heavy diesel 2.1 2.9 3.2 wt-% total diesel 12.6 15.0 24.7 wt-%VGO .2 .2 .2

These tests clearly demonstrated that the Ga-ZSM-5 catalyst can resultin significant hydroisomerization and hydrocracking of the wax, suchthat the product following this finishing step, undertaken after the FTsynthesis reaction, can be blended with crude oil and transported. Theuse of a separate finishing reactor to convert wax is superior to otherproposed options to date, including the use of a wax conversion catalystwithin the FT reactor.

Example 5 Improvement in FT Product Quality, Due to Finishing Stage

A material balanced “baseline FT” process was evaluated against the sameprocess, but with the added finishing step for the hydroisomerizationand hydrocracking of the wax produced in FT, according to informationobtained from Example 4 above. The baseline FT process utilized acatalyst containing 20 wt-% cobalt on an alumina support, and thisprocess was conducted for a sufficiently long period to establishoperational equilibrium, particularly with respect to the wax formationrate. A finishing reactor containing the Ga-ZSM-5 finishing catalyst asdescribed in Example 4 was added downstream of the baseline FT process,to evaluate its ability to convert the FT wax produced in the baselineFT process and thereby improve overall product quality, relative to theuse of the baseline FT process alone. This improvement is illustrated inTable 3 below.

TABLE 3 Improvement in FT Product Quality, Resulting from Wax Conversion(Finishing) Baseline FT plus Wax FT Conversion FT synthesis reactiontemperature, ° C. 216 216 pressure, MPa 2.07 2.07 finishing reactiontemperature, ° C. N/A 260 wt-% material recovery 96 100 wt-% carbonrecovery 95 96 % CO conversion 56 53 % C selectivity to C₁-C₃hydrocarbons 36 40 % C selectivity to C₄ ⁺ liquid hydrocarbons 26 60 % Cselectivity to wax 39 0

In view of these results, it can be seen that the combined FT synthesisand finishing stages result in the production of no wax, i.e., nohydrocarbons having melting points above room temperature. Also, byadding the finishing stage with the Ga-ZSM-5 catalyst, the selectivityto hydrocarbons useful for liquid fuels (such as C₄-C₁₉ liquidhydrocarbons), i.e., the percentage of carbon in CO converted by FTsynthesis that resulted in these hydrocarbons, was increased. Theselectivity to C₁-C₃ gaseous hydrocarbons was also slightly increased,as a result of cracking reactions that generated these products.Although these tests were not optimized in terms of minimizing the C₁-C₃gaseous hydrocarbon yield and maximizing the liquid hydrocarbon fuelyield, they nonetheless demonstrated that the use of the finishing(hydroisomerization and hydrocracking) reactions can convert essentiallyall of the wax to condensable liquid hydrocarbons useful as fuels,without an excessive generation of gaseous hydrocarbons. The completeconversion of wax was confirmed by gas chromatography-mass spectrometryanalysis (GC-MS) of the finishing product obtained after the finishingreaction.

Example 6

Integration with Biomass Hydropyrolysis to Improve Biogenic Liquid FuelYield

A comparison was made between the costs and performance of thehydropyrolysis process depicted in FIG. 6 and the process in which anintegrated CSR-FT process is added, as depicted in FIG. 5 , to increasethe yield of biogenic liquid fuels from a biomass-containing feedstock(wood). The evaluation of each case was based on a 500 ton per day (t/d)production rate of liquid fuels, for calculation purposes. Thiscomparison is provided in Table 4 below.

TABLE 4 Advantage of CSR-FT Integration with Hydropyrolysis Hydro-Hydropyrolysis, pyrolysis Integrated alone with CSR-FT Liquid fuelyield, based on biomass, wt-% 26 38 Natural gas input, based on biomass,wt-% 0 14 Capital cost estimate, millions $ 179 227 Utilities, megawatt2.0 2.0 Makeup water, liters/sec 17.9 17.9 Wastewater out, liters/sec7.1 7.1

It can be seen from this comparison that the addition of an integratedCSR-FT process, to produce additional hydrocarbons from thehydropyrolysis gaseous mixture 4 as shown in FIG. 5 , provides asubstantial improvement in the yield of these hydrocarbons (38 wt-% vs.26 wt-%, based on biomass). The carbon in these additional hydrocarbonsis derived from biomass, such that all liquid fuel from each case aboveis biogenic. It is estimated that the addition of a CSR-FT process canincrease the production rate of gasoline and diesel boiling-rangehydrocarbons from 86 gallons per ton of wood biomass to 120 gallons perton.

Overall, aspects of the invention relate to the use of dry reforming orCO₂-steam reforming to achieve high conversion of methane and/or otherhydrocarbon(s) and produce a synthesis gas product having desiredcharacteristics, including molar H₂:CO ratios as described herein.Further aspects relate to such reforming processes that use an activereforming catalyst with the ability to convert methane and/or otherhydrocarbon(s) in the presence of CO₂, or both CO₂ and H₂O, with littlecoke deposition and high catalyst stability, even in the case of feedscomprising sulfur-bearing contaminants and/or reactive compounds such asaromatic and/or olefinic hydrocarbons, with such contaminants andcompounds being associated with rapid deactivation in conventionalcatalyst systems. Yet further aspects relate to such reforming processesthat also provide a straightforward approach for direct use with furtherprocessing stages, such as Fischer-Tropsch synthesis for the productionof liquid (C₄) hydrocarbons and/or alcohols, alcohol synthesis viafermentation, or hydrogen production. Advantageously, the processes canutilize existing CO₂ present in sources of both renewable andnon-renewable methane, preferably without the removal of this CO₂,and/or can utilize lower levels of water compared to conventional steamreforming of methane. In addition, the sulfur tolerance of the reformingcatalyst is further evidenced by its activity for convertingsulfur-bearing contaminants into SO₂ and H₂S that are easily manageddownstream, if necessary, using a single acid gas removal step. Yetfurther aspects relate the integration of CO₂-steam reforming withFischer-Tropsch synthesis, as described above, optionally with afinishing stage. Those having skill in the art, with the knowledgegained from the present disclosure, will recognize that various changescan be made to these processes in attaining these and other advantages,without departing from the scope of the present disclosure. As such, itshould be understood that the features of the disclosure are susceptibleto modifications and/or substitutions without departing from the scopeof this disclosure. The specific embodiments illustrated and describedherein are for illustrative purposes only, and not limiting of theinvention as set forth in the appended claims.

1. A process for producing C₄ ⁺ hydrocarbons, the process comprising:(a) in a reforming stage, contacting a gaseous mixture comprisingmethane and an oxidant with a reforming catalyst to produce a synthesisgas product; (b) converting H₂ and CO in the synthesis gas product tohydrocarbons, including the C₄ ⁺ hydrocarbons, provided in aFischer-Tropsch (FT) product of an FT reactor; and (c) feeding the FTproduct, including the C₄ ⁺ hydrocarbons, to a finishing reactor toconvert at least about 75% of normal C₂₀ ⁺ hydrocarbons in a waxfraction of the FT product to C₄-C₁₉ hydrocarbons.
 2. The process ofclaim 1, wherein the oxidant comprises CO₂ or H₂O.
 3. The process ofclaim 2, wherein the oxidant comprises a mixture of both CO₂ and H₂O. 4.The process of claim 1, wherein step (b) is carried out with an FT feedhaving a substantially same molar H₂:CO ratio as in the synthesis gasproduct, produced in step (a).
 5. The process of claim 1, wherein, priorto step (b), water is condensed from the synthesis gas product, producedin step (a).
 6. The process of claim 1, wherein step (c) is carried outin the finishing reactor at a temperature from about 232° C. (450° F.)to about 399° C. (750° F.) and a gauge pressure from about 2.00 MPa (290psig) to about 3.10 MPa (450 psig).
 7. The process of claim 1, whereinthe C₄ ⁺ hydrocarbons, including the C₄-C₁₉ hydrocarbons, are providedin a hydroisomerization/hydrocracking product of the finishing reactor,said hydroisomerization/hydrocracking product comprising less than about1 wt-% hydrocarbons that are solid at room temperature.
 8. The processof claim 1, wherein the finishing reactor comprises a dewaxing catalysthaving hydroisomerization and/or hydrocracking activity with respect tosaid normal C₂₀ ⁺ hydrocarbons.
 9. The process of claim 8, wherein thedewaxing catalyst comprises a dewaxing active metal deposited on a solidacidic support.
 10. The process of claim 9, wherein the dewaxing activemetal is selected from Group 13 or Group 14 of the Periodic Table. 11.The process of claim 10, wherein the dewaxing active metal is gallium.12. A process for producing C₄ ⁺ hydrocarbons, the process comprising:(a) in a reforming stage, contacting a gaseous mixture comprisingmethane and an oxidant with a reforming catalyst to produce a synthesisgas product; (b) converting H₂ and CO in the synthesis gas product tohydrocarbons, including the C₄ ⁺ hydrocarbons, provided in aFischer-Tropsch (FT) product of an FT reactor; and (c) contacting afinishing feed, comprising all or a portion of the FT product, with adewaxing catalyst in a finishing reactor to convert normal C₂₀ ⁺hydrocarbons in a wax fraction of the FT product to C₄-C₁₉ hydrocarbonsby hydroisomerization and/or hydrocracking reactions; wherein thehydroisomerization and/or hydrocracking reactions consume hydrogen thatis unconverted in the FT reactor and present in the finishing feed. 13.The process of claim 12, wherein the hydrogen that is unconverted in theFT reactor, is present in the finishing feed at a concentration of atleast about 20 mol-%.
 14. The process of claim 12, wherein nosupplemental hydrogen source is added to the finishing reactor.
 15. Theprocess of claim 12, wherein a supplemental hydrogen source is added tothe finishing reactor.
 16. A process for producing C₄ ⁺ hydrocarbons,the process comprising: (a) in a reforming stage, contacting a gaseousmixture comprising methane and an oxidant with a reforming catalyst toproduce a synthesis gas product; (b) converting H₂ and CO in thesynthesis gas product to hydrocarbons, including the C₄ ⁺ hydrocarbons,provided in a Fischer-Tropsch (FT) product of an FT reactor; and (c)contacting a finishing feed, comprising all or a portion of the FTproduct, with a dewaxing catalyst in a finishing reactor to convertnormal C₂₀ ⁺ hydrocarbons in a wax fraction of the FT product to C₄-C₁₉hydrocarbons by hydroisomerization and/or hydrocracking reactions;wherein the finishing feed comprises CO that is unconverted in the FTreactor.
 17. The process of claim 16, wherein step (c) provides afinishing product exiting the finishing reactor and the process furthercomprises separating the finishing product to provide (i) C₁-C₃hydrocarbons and residual H₂, CO, and CO₂, and (ii) liquid fuelscomprising said C₄-C₁₉ hydrocarbons.
 18. The process of claim 17,further comprising separating the liquid fuels into fractions comprisinggasoline boiling-range hydrocarbons and diesel boiling-rangehydrocarbons.
 19. The process of claim 16, wherein FT reactionconditions in the FT reactor include a temperature from about 193° C.(380° F.) to about 260° C. (500° F.), a gauge pressure from about 689kPa (100 psig) to about 3.44 MPa (500 psig).
 20. The process of claim17, wherein, in step (b), a conversion of said CO in the FT reactor isat least about 30%.